Method for high efficiency reverse osmosis operation

ABSTRACT

A process for treatment of water. Hardness and non-hydroxide alkalinity are removed from feedwaters to an extent sufficient to avoid scaling when concentrated. Sparingly ionizable components in the feedwater are urged toward increased ionization by increasing the pH of the feedwater. In this manner, species such as silica become highly ionized, and (a) their rejection by membranes used in the process is significantly increased, and (b) their solubility in the reject stream from the membrane process is significantly increased. Sparingly ionized species such as boron, silica, and TOC are highly rejected. Recovery ratios of ninety percent (90%) or higher are achievable with many feedwaters, while simultaneously achieving a substantial reduction in cleaning frequency of membranes used in the process.

RELATED PATENT APPLICATIONS

This application is a continuation of prior and now pending U.S. patentapplication Ser. No. 11/342,135 filed Jan. 26, 2006, which applicationis a continuation application of prior U.S. patent application Ser. No.09/243,237 filed Feb. 2, 1999 (now abandoned), which application claimedpriority under 35 USC §121 and was a divisional of U.S. application Ser.No. 08/909,861 filed Aug. 12, 1997 (now U.S. Pat. No. 5,925,255, issuedJul. 20, 1999), which application claimed priority from prior U.S.Provisional Patent Application Ser. No. 60/077,189, filed Mar. 1, 1997,and from prior U.S. Provisional Patent Application Ser. No. 60/036,682filed Aug. 12, 1996 (converted to a Provisional application fromNon-Provisional U.S. application Ser. No. 08/695,615). The disclosuresof each of the above identified patent applications are incorporatedherein by this reference.

TECHNICAL FIELD

My invention relates to a method for the treatment of water in membranebased water treatment, purification, and concentration systems, and toapparatus for carrying out the method. In one embodiment, my inventionrelates to methods for feedwater pretreatment and for operation ofreverse osmosis (“RO”) equipment, which achieve increased soluterejection, thereby producing very high purity (low solute containing)product water, while significantly increasing on the on-streamavailability of the water treatment equipment.

BACKGROUND

A continuing demand exists for a simple, efficient and inexpensiveprocess which can reliably provide water of a desired purity, inequipment which requires a minimum of maintenance. In particular, itwould be desirable to improve efficiency of feed water usage, and lowerboth operating costs and capital costs for high purity water systems, asis required in various industries, such as semiconductors,pharmaceuticals, biotechnology, steam-electric power plants, and nuclearpower plant operations.

In most water treatment systems for the aforementioned industries, theplant design and operational parameters generally are tied to finalconcentrations (usually expressed as total dissolved solids, or “TDS”)which are tolerable in selected equipment with respect to the solubilitylimits of the sparingly soluble species present. In particular, silica,calcium sulfate, and barium sulfate often limit final concentrationsachievable. In many cases, including many nuclear power plants and manyultrapure water plant operations, boron or other compounds of similarlyacting ampholytes have a relatively low rejection across membranes inconventionally operated RO systems, and may dictate design or operatinglimitations. More commonly, the presence of such compounds result insufficiently poor reverse osmosis product water, known as permeate, thatadditional post RO treatment is required to produce an acceptably purewater. In any event, to avoid scale formation and resulting decreases inmembrane thruput, as well as potential deleterious effects on membranelife, the design and operation of a membrane based water treatment plantmust recognize the possibility of silica and other types of scaleformation, and must limit water recovery rates and operational practicesaccordingly. In fact, typical RO plant experience has been that declinesin permeate flow rates, or deterioration of permeate quality, orincreasing pressure drop across the membrane, require chemical cleaningof the membrane at regular intervals. Such cleaning has beenhistorically required because of membrane scaling, particulate fouling,or biofouling, or some combination thereof. Because of the cost,inconvenience, and production losses resulting from such membranecleaning cycles, it would be advantageous to lengthen the time betweenrequired chemical cleaning events as long as possible, whilenevertheless efficiently rejecting undesirable ionic species andreliably achieving production of high purity permeate.

Since the introduction and near universal adoption of thin filmcomposite membranes in the mid to late 1980s, the improvements in ROtechnology have been evolutionary in nature. Operating pressure neededto achieve desired rejection and flux (permeate production rate per unitof membrane surface area, commonly expressed as gallons per square footof membrane per day, or liters per square meter per day) has been slowlyreduced, while average rejection of thin film composite membrane hasimproved incrementally.

Historically, brackish water RO systems have been limited in theirallowable recovery and flux rates by the scaling and fouling tendenciesof the feedwater. It would be desirable to reduce the scaling andfouling tendencies of brackish feedwater to the point where recoverylimits would be dictated by osmotic pressure, and where flux rates canbe increased substantially, compared to limits of conventional brackishwater RO systems.

From a typical end user's point of view, several areas of improvement inRO technology—chlorine tolerance being one of them—are still sought.Thin film composite membranes, at least partly due to their surfacecharge and characteristics, are relatively prone to biological andparticulate fouling. With certain feedwaters, particularly from surfacewater sources, membrane fouling and the frequent cleaning required tocombat fouling can present some arduous, costly, and time-consumingoperational challenges.

It is known that rejection of weakly ionized species, such as totalorganic carbon (“TOC”), silica, boron, and the like, is significantlylower than rejections for strongly ionized species as sodium, chloride,etc. Since the efficiency of post-RO ion exchange is largely determinedby the level of the weak anions present in the RO permeate, it would beadvantageous to remove (reject) as many weak anions as possible in theRO unit operation. In other words, by removing (rejecting) more silica(and boron) in the RO step, a higher throughput is achievable in theion-exchange unit operation that follows the RO unit.

With the exception of an RO process disclosed in U.S. Pat. No.4,574,049, issued Mar. 4, 1986 to Pittner for a Reverse Osmosis System,which reveals a double pass (product staged) RO system design, carbondioxide typically represents the largest fraction of the anion load inRO permeate. However, a multiple pass RO configuration provides verylittle benefit under conventional RO system operating conditions, sincethe carbon dioxide content of permeate stays at the same (absolute)level and represents an even bigger fraction of the anion load. Highrejection of weak anions in a single pass RO system is, therefore,considered to be another area where significant improvement is stillsought.

In addition to increasing the rejection of the weakly ionized species,the increased rejection of strongly ionized species is also desired.

Recovery rate, or volumetric efficiency, is another parameter whereimprovements in RO system performance would be advantageous. A typicalRO system operates at about 75 percent recovery, where only 75 percentof the incoming feed to RO is used beneficially, and the rest (25percent) is discharged. With water becoming both more scarce and morecostly throughout the world, increasing the maximum achievable recoveryrate in an RO system is an important goal.

Increasing the operating flux is always important to end users, asincreased flux reduces capital costs.

Simplification and cost reduction for post-RO unit operations is alsosought by end users. This is because allowable levels of impurities inultrapure water have continually decreased with the ever tighteningdesign rules in semiconductor device geometry. Thus, lower contaminantlevels in the ultrapure water system are required. As a result, the costand complexity of the post-RO system components have dramatically grownin recent years.

High purity water processing procedures and the hardware required forcarrying them out are complex and expensive. In fact, the regenerablemixed bed ion exchange system represents, by far, the most expensive(and complicated) single unit operation/process in the entire ultrapurewater treatment system. Thus, significant improvement in thecharacteristics of the RO treated water would appreciably reduce theoverall ultrapure water system cost and complexity.

I am aware of various attempts, some in high purity water treatmentapplications and some in wastewater treatment applications, in which aneffort has been made to improve the efficiency of the rejection ofcertain ions which are sparingly soluble in aqueous solution at neutralor near neutral pH. Such attempts are largely characterized byconventional hardness removal and then raising the pH with chemicaladdition. One such method is shown in U.S. Pat. No. 5,250,185, issuedOct. 5, 1993 to Tao, et al., for Reducing Aqueous Boron Concentrationswith Reverse Osmosis Membranes Operation at High pH. In a preferredembodiment, his invention provides use of a conventional zeolitesoftener followed by a weak acid cation ion-exchanger operated in sodiumform to remove divalent cations. Due to both equipment limitations andto process design considerations, his pretreatment steps are followed bythe somewhat costly and otherwise undesirable step of dosing thefeedwater with a scale inhibitor to further prevent hardness scales fromforming. Also, although his method does provide a simultaneous hardnessand alkalinity removal step, which is of benefit in many types ofapplications which are of interest to me, his method does not providefor a high efficiency in that removal step, as is evidenced by the factthat two additional downstream softening steps are required in hisprocess. Moreover, his application pertains to, and is described andclaimed with respect to oil field produced waters containing hydrocarboncompounds (containing carbon and hydrogen only, and generally notionizable), whereas in applications which are of interest to me, suchcompounds are almost totally lacking. In applications of primaryinterest to me, a variety of naturally occurring organic acid such ashumic and fulvic acids are present, particularly in surface waterspresented for treatment.

Also, a method used in high purity water applications is disclosed inJapanese KOKAI No. Sho 58-112890, Published Jun. 29, 1984 by Yokoyama,et al., for a Method of Desalination with a Reverse Osmosis MembraneUnit. His examples show reverse osmosis units utilizing a pretreatmentprocess of strong acid cation exchange resin (“SAC”) for softening inone example, and without softening in the other example. While hisprocess will work for certain feedwaters, it does not teach howoperation at higher pH levels may be employed while still avoidingscaling of RO membranes.

In order to better understand my process; it is useful to understandsome basic water chemistry principles. With respect to calcium carbonate(CaCO₃), for example, the likelihood of occurrence of precipitation onan RO membrane in the final reject zone may be predicted by use of theLangelier Index, sometimes known as the Langelier Saturation Index(LSI). See the Naico Water Handbook, copyright 1979, by McGraw-Hill.This index is generally formulated as follows:LSI=pH_(reject)−pH_(S)where pH_(S)=the pH at saturation of CaCO₃ (reject) andpH_(S)=pCa+pAlk+C

-   -   and wherein:    -   pCa=−log of Ca⁺⁺ ion concentration (moles/liter)    -   pAlk=−log of HCO₃ ⁻ ion concentration (moles/liter)    -   C=a constant based on total ionic strength and temperature of        the RO reject.

In a given RO reject water, in order to avoid carbonate scaling, it mostpreferable to keep the LSI negative, i.e. in a condition so that CaCO₃will dissolve. However, in the field, it has been found that under someconditions, with use of certain types of anti-sealant additives, an LSIof up to about +1.5 can be tolerated, without CaCO₃ scale formationresulting. In any event, at the pH of any given RO reject, pH_(S) mustbe minimized in order to avoid undesirable scale formation. To put thisinto perspective, consider that in any RO pretreatment operation, it canbe anticipated that there will always be at least some leakage ofcalcium from the softening step. Thus, depending upon the raw feedwaterhardness and the pretreatment process scheme practiced, a lower limit onthe achievable value of the pCa term, due to the concentration of theCa⁺⁺ ion present in the treated RO feedwater, can be anticipated.Furthermore, in all events, the value of C is fixed by the total ionicstrength and by the temperature. Thus, to keep the LSI in an acceptablerange—in order to provide scale free RO operation—the leakage of calcium(as well as other hardness such as magnesium) becomes a critical factor.The Tao et al. patent, identified above, approaches this problem byproviding various types of softeners in series. Specifically, he simplyaccepts the inevitably high capital and operating costs associatedtherewith. Yokoyama, on the other hand, evidently decided to limit ROoperation to a pH which is consistent with the degree of calciumremoval. When he operates with RO reject at a pH of 9, assuming 0.1 ppmof Ca⁺⁺ leakage from the ion exchange train disclosed, and aconcentration factor of 5 (“5×”) in the RO, his RO operation may beexpected to provide an RO reject with an LSI of about −0.5. That LSI isacceptable for non-scaling operation, with or without scale inhibitors.However, if the pH in Yokoyama's example were increased to 11, forexample, given the same pretreatment method, an LSI of about +2.4 mightbe expected. In such a case, the Langelier Saturation Index of thereject water would be well above the level where current anti-sealantshave the ability to provide scale free RO operation.

Thus, for the most part, the prior art methods known to me have one ormore of the following shortcomings:

-   -   (a) they do not reliably achieve the extremely low hardness and        non-hydroxide alkalinity levels necessary for essentially scale        free operation at very high pH levels;    -   (b) they rely on redundant and expensive capital equipment, with        attendant operating costs, to minimize hardness leakage;    -   (c) they depend primarily on hardness reduction to reduce the        LSI of the RO reject (and do not include provisions for high        efficiency dealkalization); and    -   (d) they rely on anti-scaling additives to prevent scale        formation.

Thus, the advantages of my simple treatment process which exploits (i)hardness removal to very low residual levels, and (ii) efficientdealkalization, to allow extended trouble free RO operation at high pHlevels, are important and self-evident.

Moreover, because of upper concentration factor limits due to thetendency of scale to form, RO systems are often unable to use abouttwenty five (25%) or more of the raw feedwater. Also, at recoverieslevels greater than approximately seventy five percent (75%) or somewhatlower, depending upon raw water chemistry, the control of chemicalscaling and biological fouling in conventional RO systems becomes almostunmanageably difficult when trying to achieve long run times. Therefore,widespread commercial use of RO systems with water recovery in excess ofabout seventy five percent (75%) has not been accomplished.

As water is becoming increasingly expensive, or in short supply, orboth, it would be desirable to increase the ratio of treated productwater to raw water feed in RO systems. Therefore, it can be appreciatedthat it would be desirable to achieve reduced costs of water treatmentby enabling water treatment at higher overall recovery rates than iscommonly achieved today. Finally, it would be clearly desirable to meetsuch increasingly difficult water treatment objectives with bettersystem availability and longer run times than is commonly achievedtoday.

In so far as I am aware, no one heretofore has thought it feasible tooperate a reverse osmosis based water treatment system at higher thanabout pH 9, in continuous, sustainable, long term operations to producea highly purified treated water product. The conventional engineeringapproach has been to design around or battle scale formation, by use ofmoderate pH, by limiting final concentration and resulting waterrecovery, by use of chemical additives. Historically, cellulose acetatemembranes were limited in operation to a pH range of roughly 4 to 7.Newer polyamide and thin-film-composite type membranes havetraditionally been operated in the pH range of roughly from about 4 toabout 8. Although higher pH operation has occasionally been attemptedfor a few special purposes, it has usually been in non-silica relatedapplications. And, although higher pH operation has been utilized insecond pass RO applications where silica was of concern, in so far as Iam aware, it has only been accomplished after a first pass RO operationwith a neutral or near neutral pH of operation. In those cases whereorganics are of specific concern, then the pH may often range to below5, and preferably, below 4.

In contrast to prior art methods for water treatment, the method taughtherein uses the essential design philosophy of virtually eliminating anypossible occurrence of scaling phenomenon during first pass operation atthe maximum feasible pH using the available membranes, while maintainingthe desired concentration factor, and taking the benefit of waterrecovery that results.

SUMMARY

I have now invented a novel water treatment method based on aggressivehardness and alkalinity removal, followed by membrane separation at highpH, to produce a high quality permeate with extremely low silicaconcentration.

In a unique feedwater treatment process, raw feedwaters of suitablechemical composition are treated with a weak acid cation ion exchangeresin, operated in the hydrogen form, to simultaneously remove hardnessand alkalinity. The weak acid cation ion exchange resins can be operatedat incoming raw feedwater hardness and alkalinity levels well abovethose that would cause conventional ion exchange systems to fail due tohardness breakthrough.

The preferred treatment train design used in my wastewater treatmentplant overcomes a number of important and serious problems. First, thelow hardness, combined with virtual elimination of non-hydroxidealkalinity, substantially eliminates the precipitation of scale formingcompounds associated with sulfate, carbonate, or silicate anions. Thus,cleaning requirements are minimized. This is important commerciallybecause it enables a water treatment plant to avoid lost waterproduction which would otherwise undesirably require increased treatmentplant size to accommodate for the lost production during cleaningcycles. Second, the preferred high pH operational conditions enable ahigh degree of ionization to be achieved in various species which aresparingly ionized at neutral or near neutral pH in aqueous solution, toenable such species to be preferentially rejected by the membranesystem. Finally, operation at high pH provides protection againstbiological contamination, thus preventing undesirable contamination ofproduct water. At the preferred high operational pH, bacteria andendotoxins are effectively destroyed. In essence, water treatmentsystems operated according to the teachings herein normally operate atconditions which might ordinarily be considered cleaning conditions forconventional RO systems.

I have now developed a novel process design for use in treatment ofwater. In one embodiment, the process involves treatment of a feedwaterstream which is characterized by the presence of (i) hardness, (ii)alkalinity, and (iii) molecular species which are sparingly ionized whenin neutral or near neutral pH aqueous solutions, to produce a low solutecontaining product stream and a high solute containing reject stream.The process involves effectively eliminating the tendency of the rawfeedwater to form scale when the raw feedwater is concentrated todesired concentration factor at a selected pH, by effecting, in anyorder, one or more of the following (i) removing hardness from the rawfeedwater stream, (ii) removing alkalinity from the raw feedwaterstream, or (iii) removing dissolved gases created during the hardnessremoval step. Then, the pH of the feedwater is raised to a selected pHof at least about 8.5, or up to 9.0, or up to about 10, or preferably(with currently available thin film composite type membranes) to a rangebetween 10 and 11, or otherwise in excess of 11, and more preferably toabout 12 or somewhat more, until the benefits gained by high rejectionrates of silica and other species is outweighed by the additional cost.With currently available thin film composite membranes, controlling thepH at up to about 10.5 provides most of the benefits of this methodwithout compromise of long-term membrane life. The pH increase isaccomplished by adding a selected base to the softened and dealkalatedfeedstream, preferably by direct injection or alternately by the use ofanion ion-exchange. The pH increase urges the molecular species whichare sparingly ionized when in neutral or near neutral pH towardincreased ionization. An alternate concept is that the protonatable,i.e., proton accepting substances, or bases, are increased. The pHadjusted feedwater is then sent through membrane separation equipment,typically of the reverse osmosis type, but alternately of nanofiltrationor other suitable type or configuration which is otherwise available, orwhich may in the future become available, and in which the currentmethod may be practiced, to produce a reject stream and a productstream. The membrane separation equipment is ideally of the type whichhas a semi-permeable membrane which substantially resists passage ofionized species therethrough. It is important that in my process, themembrane separation equipment produces a product stream which issubstantially free of the normally undesirable species which aresparingly ionized when in neutral or near neutral pH in aqueoussolutions.

OBJECTS, ADVANTAGES, AND FEATURES

From the foregoing, it will be apparent that one important and primaryobject of the present invention resides in the provision of a novelmethod for treatment of water to reliably and continuously produce overlong operational cycles a water product stream of a preselectedextremely high purity quality standard.

More specifically, an important object of my invention is to provide amembrane based water treatment method which is capable of avoidingcommon scaling and fouling problems, so as to reliably provide a methodof high purity water generation when operating at high efficiency.

Other important but more specific objects of the invention reside in theprovision of a method for water treatment as described in the precedingparagraph which:

-   -   allows the removal of hardness and alkalinity from a selected        feedwater to be done in a simple, direct manner;    -   has a minimum of unit process requirements; minimize or avoid        complex chemical feed systems;    -   requires less physical space than existing technology water        treatment plants;    -   is easy to construct, to start, and to service;    -   has high efficiency rates, that is, they provide high product        water outputs relative to the quantity of feedwater input to the        water treatment plant;    -   in conjunction with the preceding object, provide lower unit        costs to the water treatment plant operator and thus to the        water user, than is presently the case;    -   in conjunction with the just mentioned object, results in less        chemical usage than in most water treatment facilities, by        virtually eliminating use of some types of heretofore commonly        used chemical additives, particularly scale inhibitors.

A feature of one embodiment of the present invention is the use of aunique combination of weak acid cation ion-exchange with substantiallycomplete hardness and alkalinity removal, and subsequent high•pH ROoperation, thereby enabling the water treatment plant to minimize thepercentage of reject water. This results in high overall cycleefficiencies.

Another feature of the present invention is the use of a high pHoperation to highly ionize weakly ionizable species such as silica,boron, or TOC, thus enabling operation with silica, boron, or TOCrejection levels considerably exceeding the limits of conventional ROtreatment systems when treating feedwaters of comparable chemistry.

Yet another feature of the present invention is the capability toretrofit existing RO plants to operate according to the present processdesign, to increase capacity without increasing the RO membranerequirements.

Another feature of the present invention is the ability to providehigher purity product water while operating at higher flux levels thanhas heretofore been feasible with conventional RO system designs.

Other important objects, features, and additional advantages of myinvention will become apparent to those skilled in the art from theforegoing, and from the detailed description which follows, and from theappended claims, in conjunction with the accompanying drawing.

BRIEF DESCRIPTION OF THE DRAWING

In the drawing, identical features shown in the several figures will bereferred to by identical reference numerals without further mention.

FIG. 1 illustrates the percentage ionization of silica ions in aqueoussolution as a function of pH.

FIG. 2 illustrates a first embodiment of my method for high efficiencyreverse osmosis operation, showing use of a weak acid cation exchangeunit for simultaneous hardness and non-hydroxide alkalinity removal.

FIG. 3 shows a second embodiment of my method for high efficiencyreverse osmosis operation, wherein hardness is reduced by sodium zeolitesoftening and optional lime or lime/soda softening.

FIG. 4 shows a third embodiment of my method for high efficiency reverseosmosis operation, showing the equipment configuration where alkalinityin raw feedwater can be efficiently and adequately reduced by acidaddition, and where hardness may optionally be reduced by lime orlime/soda softening.

FIG. 5 illustrates the differential pressure, in pounds per square inchversus time (PSID v. Months) for a reverse osmosis membrane employed inpilot reverse osmosis test equipment utilizing my novel process.

FIG. 6 illustrates the normalized permeate flow, in liters per minuteversus time, for a reverse osmosis membrane employed in pilot reverseosmosis test equipment utilizing my novel process.

FIG. 7 illustrates the silica concentration in the reverse osmosisreject stream in pilot reverse osmosis test equipment utilizing my novelprocess.

FIG. 8 illustrates the rejection percentage of silica versus time, for areverse osmosis membrane employed in pilot testing of my novel process.

FIG. 9 describes the use of my method of RO system operation when usinga multipass RO process to sequentially process a portion of initialfeedwater to produce a permeate which has been passed through more thanone RO membrane.

FIG. 10 illustrates the use of my method of RO system operation forboiler feed makeup water, or for cooling tower makeup water, or forscrubber makeup water.

FIG. 11 illustrates the use of my method of RO operation in combinationwith continuous electrodeionization equipment for high purity waterproduction.

FIG. 12 illustrates a process flow diagram for one configuration of myhigh efficiency RO process.

FIG. 13 illustrates a system schematic for a conventional RO systemprocess design.

FIG. 14 illustrates an exemplary process flow diagram for my highefficiency RO process, utilizing the design and operational conceptstaught herein.

DETAILED DESCRIPTION

I have developed a new method for process design and operation of ROsystems. This new method for process design and operation of RO systemshas been thoroughly tested. The process has shown that it is capable ofachieving important improvements in RO operational objectives.

Attributes which characterize my HERO™ brand RO process design andoperation include:

-   -   (1) Very high rejection of all contaminants, especially weak        acid anions such as TOC, silica, boron, etc.    -   (2) Very high achievable recovery—ninety percent (90%) or higher        recovery can be achieved.    -   (3) Biological fouling is essentially eliminated.    -   (4) Particulate fouling is substantially reduced.    -   (5) Cleaning frequency is substantially reduced.    -   (6) Removal of chlorine from the feedwater may not be needed,        due to the resulting chemical species present at the high        operating pH, or in some cases, by eliminating the need to add        chlorine in the first place.    -   (7) Addition of scale inhibitors is virtually eliminated.    -   (8) Substantially higher flux is achieved.    -   (9) Reduced overall capital cost, compared to conventional RO        systems.    -   (10) Reduced overall operating cost, compared to conventional RO        systems.    -   (11) The complexity of an ultrapure water system is        significantly reduced.

The HERO brand RO system is highly site-specific. Individual processsteps are customized to fit the specific feedwater at a specific site.Regardless of the difference in pretreatment process for differentsites, one process parameter is common for all applications, namely thatthe RO system is operated at the highest feasible reject pH. Consistentwith the highest allowable pH limit for currently available RO membranes(for example, pH 11.0 for FILMTEC(R) brand RO elements), a typical HERObrand RO system is designed to operate at pH of up to approximately 11,as measured in the RO reject stream.

Because of the very high concentration factors (i.e. percent recovery)allowed by my HERO brand RO process, the RO feed pH is correspondinglylower. For example, in a system operating at ninety percent (90%)recovery, a feed pH of 10.0 will produce a reject stream at anapproximate pH of 11, provided that the RO feed is only slightlybuffered by the presence of carbonate, phosphate, etc. Unlikeconventional RO systems, typically operated at about seventy fivepercent (75%) recovery, a HERO brand RO system can be routinely operatedat ninety percent (90%) or greater recovery, limited only by osmoticpressure of the RO reject. The pH increase from RO feed to reject ismagnified at very high recoveries. Thus, the maximum allowable pH isspecifically applicable for the RO reject conditions.

In order to operate an RO system with reject up to near pH 11, or atabout pH 11, or above, several process conditions must be met in orderto effectively eliminate the potential for scale formation on the ROmembrane. Some of those process conditions are also necessary foroperating an RO system at very high recovery rate. Such processconditions are as follows:

-   -   (1) Calcium, magnesium, strontium, and barium concentration in        the RO feed must be substantially eliminated, preferably to near        zero, and most preferably, to essentially zero.    -   (2) Aluminum, iron, and manganese content including organically        bound species, as well as the presence of colloidal particles        containing such materials, should be substantially eliminated,        and preferably to near zero.    -   (3) Buffering anions (specifically bicarbonate, or carbonate,        and/or phosphate species) should be reduced to as low of a level        as can be practically achieved.

The selection of specific operations and control points to fulfill theabove process condition requirements is influenced by thecharacteristics of each specific feedwater. The percent recovery needed(or desired for a specific application) also affects the operations andcontrol point criteria as well. FIG. 2 represents a highly costeffective RO unit process sequence.

The first step is to adjust the hardness-to-alkalinity ratio of thefeedwater, if needed. Optimizing this ratio, which is typically done byalkali addition, makes complete hardness removal feasible in the nextprocess step.

The second step in the RO process train involves the utilization of aweak acid cation (WAC) resin (e.g. DOWEX® MAC-3, or Lewatit CNP-80,Amberlite® IRC-86). Operated in hydrogen form, the WAC resin removeshardness quantitatively, given the proper hardness-to-alkalinity ratioof the influent. The hydrogen ions liberated in the cation exchangeprocess react with the alkalinity and produce carbonic acid (H₂CO₃)₂,which is dissolved in the WAC effluent.

The third step involves adding acid to the WAC effluent to destroy theremaining alkalinity, if any such alkalinity is present. Totalalkalinity removal at this step is important in order to achieve veryhigh recovery across the RO system.

In a fourth step, the acidified effluent, containing virtually zerohardness and alkalinity, is then treated for carbon dioxide removal.This removal can be accomplished in a forced/induced draft decarbonatoror in an existing vacuum degasifier of either packed bed or gaspermeable membrane barrier design. The decarbonated, essentially zerohardness, essentially zero alkalinity water, is then injected with asoluble alkali, preferably for adjusting pH to 10.0 or higher, and mostpreferably to the pH as needed to achieve pH up to at or near 11.0 inthe RO reject.

The next step consists of operating the RO system in such a manner thatthe pH of the reject is approximately, but preferably not appreciablyhigher than, 11.0. Note that this pH 11 limitation is applicable simplywith respect to currently available RO membranes. An exemplary membrane,with the highest pH tolerance capability, is a FILMTEC type FT30membrane. If RO membranes with a higher pH tolerance capability becomeavailable in the future, then the maximum allowable RO reject pH can beraised accordingly, with concomitant benefits from the higher pH, inexcess of 11.0.

Feedwaters utilized for production of high purity water, as well asthose encountered in wastewater treatment, include the presence ofsilicon dioxide (also known as silica or SiO₂) in one form or another,depending upon pH and the other species present in the water. Formembrane separation systems, and in particular for RO type membraneseparation systems, scaling of the membrane due to silica is to bereligiously avoided. This is because (a) silica forms relatively hardscale that reduces productivity of the membrane, (b) is usually ratherdifficult to remove, (c) the scale removal process produces undesirablequantities of spent cleaning chemicals, and (d) cleaning cycles resultin undesirable and unproductive off-line periods for the equipment.Therefore, regardless of the level of silica in the incoming rawfeedwater, operation of conventional membrane separation processesgenerally involves concentration of SiO₂ in the high total dissolvedsolids (“TDS”) stream to a level not appreciably in excess of 150 ppm ofSiO₂ (as SiO₂). Typically, RO systems are operated at lowered recoveryrates, where necessary, to prevent silica concentration in the rejectstream from exceeding roughly 150 ppm.

Scaling due to various scale forming compounds, such as calcium sulfate,calcium carbonate, and the like, can be predicted by those of ordinaryskill in the art and to whom this specification is directed, by use ofthe Langlier Saturation Index, as discussed above, or other availablesolubility data. Operating parameters, including temperature, pH,permeate and reject flow rates, must be properly accounted for, as wellas the various species of ions in the raw feedwater, and those speciesadded during pretreatment.

I have found that by reliable hardness and non-hydroxide alkalinityremoval, to levels which effectively avoid formation of scale at aselected pH for RO operation, the concentration of SiO₂ in the RO rejectstream can be safely increased to 450 ppm or more. This is accomplishedby increasing the pH of the feedwater to the RO system, and without useof scale-inhibition chemicals. Moreover, even with this increase ofsilica in the RO reject, the level of silica contamination in the ROpermeate is preferentially and substantially decreased, when compared tothe silica which might be anticipated under conventional RO processconditions.

It is commonly understood that the solubility of silica increases withincreasing pH, and that silica is quite soluble in high pH aqueoussolution. Along with solubility, the degree of ionization of silica alsoincreases with increasing pH. While the increase in silica solubility isnot directly proportional to the degree of ionization, the rate ofincrease in silica solubility is basically proportional to the rate ofchange in ionization. This discrepancy between solubility and ionizationis explained by the fact that even undissociated silica exhibits somesolubility in aqueous solutions, typically up to about one hundredtwenty (120) ppm to one hundred sixty (160) ppm, depending upontemperature and other factors. In comparison, silica solubility at pH 11is in excess of one thousand five hundred (1,500) ppm at ambienttemperature; silica is increasingly soluble as temperature and/or pHincreases.

Silica is very weakly ionized when in neutral or near neutral aqueoussolutions and is generally considered to exist as undissociated(meta/ortho-) silicic acid (H₄SiO₄) in most naturally occurring waterswith a pH of up to about 8. The dissociation constant (pKa) value forthe first stage of dissociation of silica has been reported atapproximately 9.7, which indicates that silica is approximately fiftypercent (50%) ionized at a pH of 9.7; the other fifty percent (50%)remains as undissociated (ortho) silicic acid at that pH. A graphicalrepresentation of the relationship between pH and the percent silicaionization is shown in FIG. 1. Clearly, it would be advantageous, wheresilica ionization is desired, to operate at a pH in excess of 10, andmore preferably, in excess of 11, and yet more preferably, in excess of12.

The understanding of silica ionization phenomenon is important since therejection of most species across the membranes of membrane separationequipment is enhanced by increased ionization. Consequently, silicarejection by an RO membrane can be expected to improve as the degree ofionization increases; with respect to silica, ionization increases athigher pH. Therefore, increasing the pH of the RO operation thusprovides major benefits. First, silica solubility can be radicallyincreased, even while remaining within the current pH limitations ofexisting commercial thin film composite type RO membranes. Second,silica rejection is increased significantly at high pH levels,corresponding to the increased degree of ionization of the silica.

To gain maximum benefit from silica ionization at high pH, the RO systemshould be operated at a pH as high as possible, given the limitationsimposed by membrane chemistry and by the membrane manufacturer'swarranty. Preferably, the RO system is operated at a pH of about 10 orabove, and more preferably, at 10.5 or above, and most preferably, at apH of 11 or higher. This contrasts with typical RO operation practice,where operating pH has been limited to about 8.5, in order to avoidscale formation, particularly silica and carbonate scales.

Referring again to FIG. 2, one embodiment of my process for membraneseparation equipment operation is shown. In this method, raw water 10 isfirst treated in a weak acid cation ion exchange unit 12, where hardnessand bicarbonate alkalinity are simultaneously removed. For those caseswhere raw water 10 hardness is greater than alkalinity, operation of theweak acid cation ion exchange unit 12 must be facilitated by addition ofa source of alkalinity 13, such as by addition of an aqueous solution ofsodium carbonate (Na₂CO₃). Preferably, the WAC unit 12 is operated inthe hydrogen form for ease of operation and regeneration. However, itwould also work in the sodium form, followed by acid addition. In anycase, in the just mentioned case and otherwise optionally whereappropriate, acid 14 is added by pump 16 to the effluent 18 from the WACunit(s) 12 to enhance bicarbonate destruction. Then, the carbon dioxide19 which has been created in the WAC (and/or by acid addition) isremoved, preferably in an atmospheric pressure or vacuum degassifier 20.Finally, an alkali 22 (base) is added, preferably by pump 24 injectionof liquid solution, to increase the pH of the feedwater 25 to a selectedlevel. Any of a variety of conveniently available and cost effectivebase products may be used, provided that no appreciable scaling tendencyis introduced. Besides use of common sodium hydroxide, other chemicalssuch as sodium carbonate, potassium hydroxide, or potassium carbonatemight be selected. In fact, in certain cases, an organic base, such as apyridine type compound, may be used effectively to carry out thisprocess. In any event, pressurization of feedwater 25 for the membraneprocess is accomplished by high pressure pump 26 before transfer to theRO type membrane separation unit 30 as shown. Alternately, alkali (base)addition to the feedwater may be accomplished by passing the feedwaterthrough an anion ion-exchange unit 31 to increase the pH to a desiredlevel. The pH of the feedwater is raised to a selected pH of at leastabout 8.5 or 9.0, or up to about 10, or preferably (with currentlyavailable thin film composite type membranes) to a range between 10 and11, or otherwise in excess of 11, and more preferably to 12 or more, andmost preferably, to 13 or more. With currently available thin filmcomposite type RO membranes, such as those sold by DOW CHEMICAL ofMidland, Mich. under their FILMTECH brand by their FILMTEC, INC.subsidiary, controlling the pH to about 10.5 provides most of thebenefits of this method without compromise of long-term membrane life.However, to increase silica solubility, and silica rejection, membranesallowing the pH to be raised to at least about 11, or more preferably toat least about 12, or most preferably, to at least about 13, would bedesirable. Thus, it can be appreciated that my method may be used toeven further advantage when membranes with long life expectancy at suchelevated pH's become commercially available.

Reject 32 from membrane separation unit 30 may be sewered or sent tofurther treatment, as appropriate in particular site circumstances.Permeate 34 from membrane separation unit 30 may utilized “as is” or maybe further purified to remove residual contamination, for example, forhigh purity water users such as semiconductor manufacturing, where 18.2meg ohm purity water is desired. A conventional post-RO treatment trainfor production of high purity water 38 in the semiconductor industryincludes a cation exchanger 40, followed by an anion exchanger 42, withprimary 44 and secondary 46 mixed bed polisher ion exchange units.Somewhat different post RO treatment trains may be utilized to meet theparticularized needs of a given site, raw water chemistry, and end use,without departing from the advantages and benefits which may be gainedby the RO process method disclosed herein. For example, it may bedesirable in some circumstances to omit the cation 40 and anion 42ion-exchangers, and bypass the RO permeate via line 47 to directly reachthe primary mixed bed 44 and polish mixed bed 46 ion-exchange units.Finally, in many ultrapure water plants, the product from the polishingmixed bed ion-exchange units 46 is currently further treated in finalfiltration units 48 and ultraviolet irradiation units 49 to eliminateparticulates and biofouling, respectively. Additional treatmentoperations may added as appropriate to meet the needs of a particularend user.

Another distinct and unique advantage of my method of RO systemoperation is that it may be possible, under various raw feedwaterchemistry and operating conditions, to operate the entire post-RO ionexchange train (i.e., ion-exchange, units 40, 42, 44, and 46) withoutregeneration. Depending upon chemistry, it may be possible to simplyreplace the cation 40 and anion 42 exchangers. In the more usual case,the secondary or polishing mixed bed unit 46 may be replaced with newresin, and the old polishing resin moved to the primary bed 44 position.This is possible, particularly in ultrapure and boiler feed type watertreatment systems, because the polishing mixed bed unit 46 is controlledby ending operation when the silica, boron, or other ion leakage reachesa predetermined value. When the predetermined ion leakage value isreached, the then polishing mixed bed unit 46 is substituted for, andplaced into the position of, the primary mixed bed ion-exchange unit 44.When the change over of mixed bed ion-exchange units is made, the “old”primary mixed bed unit 44 resin is taken out, and either discarded orsold to other less demanding resin users. New resin is then put into the“old” primary mixed bed ion-exchange unit 44, whereupon it becomes the“new” polishing mixed bed ion exchange unit 46.

In other embodiments, and as suited to meet the particularized needs ofa selected raw feedwater chemistry, various forms of hardness removalmay be utilized, including sodium form strong acid cation exchange 50,followed by acidification (see FIG. 3) or even the use of a lime 52 (orsimilar lime/soda) softener as an additional pretreatment step to eithersodium form strong acid cation exchange 50 or weak acid cation exchange12 (see FIGS. 2 and 3).

For particularly soft waters, the lime or lime/soda softener 52 may betotally inappropriate, and this method may proceed with no softening ofthe raw water, and only a simple acid 14 feed before decarbonization, ascan be seen in FIG. 4. On the other hand, where softening isappropriate, some raw feedwaters can be appropriately treated forreductions in hardness and alkalinity to a desired extent by softener52. Regardless of the equipment configuration selected for treatment ofa particular raw water chemistry, the key process parameters are (a) toremove those cations which, in combination with other species present athigh pH, would tend to precipitate sparingly soluble salts on themembrane surfaces, and (b) eliminate non-hydroxide alkalinity to themaximum extent feasible, to further protect against precipitation ofscales on the membrane surfaces.

The weak acid cation (“WAC”) ion-exchange resins used in the first stepof the preferred embodiment of my method, as illustrated in FIG. 2, arequite efficient in the removal of hardness associated with alkalinity.Such a reaction proceeds as follows:Ca⁺⁺+2RCOOH--->(RCOO)₂Ca+2H⁺

Then, the hydrogen combines with the bicarbonate to form carbonic acid,which when depressurized, forms water and carbon dioxide, as follows:H⁺+HCO₃ ⁻---->H₂CO₃---->H₂O+CO₂

Regeneration of the resin is accomplished by use of convenientlyavailable and cost effective acid. It is well known by those in the artthat regeneration of WAC ion-exchange resins may proceed quiteefficiently, at near stoichiometric levels (generally, not more thanabout one hundred and twenty percent (120%) of ideal levels).Preferably, hydrochloric acid may be used, since in such cases highlysoluble calcium chloride would be produced, and the regeneration processwould not pose the potential danger of formation of insoluble sulfateprecipitates, such as calcium sulfate, even with high strength acids.However, by use of a staged regeneration procedures, i.e., by using alow concentration acid followed by a higher concentration acid, it ispossible to reliably utilize other acids, including sulfuric acid(H₂SO₄), while still avoiding undesirable precipitates on the resin. Inthis manner, hardness ions are solubilized to form soluble salts, whichare eluted from the resin bed and are typically sewered. Use of sulfuricacid is particularly advantageous in semiconductor manufacturingoperations, since such plants typically use large quantities of suchacid, and waste or spent acid may be advantageously utilized forregeneration of a weak acid cation exchange bed.

Other polyvalent cations, most commonly iron (Fe⁺⁺/Fe⁺⁺⁺), magnesium(Mg⁺⁺), barium (Ba⁺⁺), strontium (Sr⁺⁺), aluminum (Al⁺⁺⁺), and manganese(Mn⁺⁺/Mn⁺⁺⁺⁺), are also removed by the WAC resin. Since the presence ofeven very small quantities of hardness or other species of decreasingsolubility at increasing pH will result in precipitation of sparinglysoluble salts under the process conditions present in my process, caremust be taken to prevent precipitation on the membrane of the substancessuch as of calcium carbonate, calcium hydroxide, magnesium hydroxide,and magnesium silicate. One precaution that should be observed is thatboth hardness and non-hydroxide forms of alkalinity should beaggressively reduced in the feedwater, prior to upward pH adjustment toselected RO operating conditions. Once hardness and non-hydroxide formsof alkalinity have been removed, then the desired pH increase may beaccomplished with any convenient alkali source, such as sodium orpotassium alkali, or by anion exchange. Once this pretreatment has beenthoroughly accomplished, then an RO system can be safely operated atvery high pH levels, in order to take advantage of the aforementionedsilica solubility.

In cases where raw water composition is such that sodium zeolitesoftening is advantageous, as is depicted in FIG. 3, elimination ofcalcium hardness proceeds as follows:Ca⁺²+Na₂X--->CaX+2Na⁺

Then, bicarbonate alkalinity is converted to carbon dioxide, with aselected acid, in a manner similar to the following:NaHC0₃+HCl--->NaCl+H₂O+CO₂

For those waters where lime softening may be an acceptable or preferredmethod for initial hardness and alkalinity reduction, the addition oflime to the water reduces calcium and magnesium hardness, and associatedbicarbonate alkalinity, as follows:Ca(HCO₃)₂+Ca(OH)₂--->2CaCO₃↓+2H₂OMg(HCO₃)₂+2Ca(OH)₂--->Mg(OH)₂↓+2CaCO₃+2H₂O

This process configuration is depicted as an alternate embodiment of mymethod, as illustrated in FIGS. 3 and 4. In the cases where lime orlime/soda softening is used, however, extreme care must be used inevaluating the performance of the remainder of the pro-treatment system,since the solubility of hardness ions remains appreciable in thesoftener 52 effluent stream 54.

For most feedwaters, particularly where lime or lime/soda softening isnot employed, the use of a carbon dioxide removal step significantlyenhances cost-effectiveness of the process when carried out prior to thepH increase. This also helps to maintain a lower total alkalinity levelin the feed to the RO, thus providing a greater margin of safety againstscaling due to hardness leakage from the cation removal step.Dealkalization by carbon dioxide removal also helps to enhance silicarejection, due to the lack of competing species. This is because therejection of one weakly ionized anion is affected by the presence andconcentration of other weakly ionized anions in the feedwater; thisapplies to weakly ionized anions such as boron, organic acids (TOC),cyanide, fluoride, and certain arsenic and selenium compounds.

Since the high pH operation also increases ionization of other weaklyionized anions, including borate, organic acids (TOC), cyanide,fluoride, and certain arsenic and selenium compounds, their rejectionrates are enhanced in an RO membrane system. Consequently, in general,my method may be advantageously applied to reject across the membranemost weak acids with a pKa₁ of about 7.5 or higher. Silica rejection canbe increased to about 99.95%, or more, from a conventional baseline ofabout 99% rejection; this amounts to at least one order of magnitudedecrease in the amount of silica escaping into the permeate, thusproviding a ten plus (10⁺) fold increase in running life for the silicascavenging ion-exchange resin bed, namely anion exchanger 42 and themixed bed units.

In the case of cyanide, rejections in a first pass RO of in excess ofninety percent (90%) can be attained, in contrast with a more typicalrange of about fifty percent (50%) or so with conventional RO processes.

Similar to the case for silica, boron rejection can be increased, from aconventional baseline from a range of about 60-70% to 99% and higher, byoperation at a suitably high pH. The beneficial effects on rejectionpercentage due to higher pH operation start at a slightly lower pH inthe case of boron, since the pKa for boron is 9.14, roughly one-half pHunit higher than that for orthosilic acid, namely 9.7. The beneficialeffects of high pH operation are much more pronounced in the case ofboron, however, because orthosilic acid (H₂SiO₄) in aqueous solutiontypically includes six molecules of water of hydration, whereas boricacid (H₃BO₃) typically has no attached hydrating water molecules. Thus,the orthosilic acid molecule is very large with respect to membrane poresize as compared to boric acid, no matter what the pH, and as a result,silica has much higher normal rejection rates. Consequently, theincreased ionization of boric acid when operating at a pH in excess ofabout 9.1 is extremely beneficial, and increasingly so as pH rises tobetween 10 and 11, or the currently preferred control point ofapproximately 10.5. The boron rejection effect would be even furtherenhanced when operating an RO system at a pH of 12 or even 13, whencommercial membranes become available for such practice.

Example Pilot Test

A pilot water treatment system was set up to test the efficacy of themethod disclosed. The pilot water treatment system was designed fortreating an incoming raw city water supply to provide high purityproduct water for potential future use in a semiconductor manufacturingplant. The objectives were (a) to increase recovery, so as to minimizewater usage, (b) to increase the purity of treated water, and (c) toincrease the average time between membrane cleanings. The pilot systemperformed a series of tests. In each of the tests, the system wasstarted up with 450 ppm or higher silica level in the RO reject. Thepilot plant system was operated continuously until either (a) a tenpercent (10%) decline in normalized RO permeate water flow wasexperienced, or (b) a fifteen percent (15%) increase in axialdifferential pressure across the RO membrane was reached. The pilot testwas performed with a membrane separation unit including a Dow/Filmtec ROMembrane Model FT30/BW4040, which was operated at pressures from about130 psig to about 185 psig, with feedwater temperatures ranging fromabout 20° C. to about 25° C., and at feedwater rates of up to about 8 USgallons per minute (30 liters per minute) maximum. As seen in FIG. 6,long term normalized permeate flows of slightly more than 5 US gallonsper minute (about 20 liters per minute) were tested. The pilot testapparatus included a pair of weak acid cation ion exchange beds operatedin parallel, utilizing Rohm and Haas Company (Philadelphia, Pa.) weakacid cation resin product number IRC-86, followed by a forced airdecarbonator, sodium hydroxide injection, separation of the treatedfeedwater by the RO membrane into a reject stream and a permeate stream.

Table 1 presents the chemical analyses of from the pilot plant operationfor raw water, RO reject, and RO permeate. The Table 1 also shows therejection rates achieved in the pilot RO operation, and compares thoserates with those achieved with a conventional RO system operating on thesame feedwater. In particular, note the level of silica in the rawfeedwater (67 ppm) and in the RO reject (480 ppm). The silicaconcentration in the RO reject is roughly three times that normallyachievable in reject water from conventional RO process configurations.Moreover, even at the high concentration of silica in the RO reject,improved rejection of silica is seen, in that silica rejection of 99.87%was achieved, compared with rejections ranging from about 95% up toabout 99% with a conventional RO system on the same feedwater.

In fact, improved rejection rates were experienced with all importantchemical species over the rejection rates experienced with conventionalRO, as is clear from the data presented in Table 1.

TABLE 1 PILOT TEST ANALYTICAL RESULTS Raw Conventional Feed Pilot ROReject Pilot RO Pilot RO RO Rejection (ppm) (ppm) permeate (ppm)Rejection (%) (%) Sodium 29.9 460 0 955 99.73 95-98 Potassium 6.4 18.7<0.003 99.98+ 90-95 Calcium 34 <0.1 <0.003 — Magnesium 5.3 <0.1 <0.0001— Chloride 12.1 78.1 <0.004 99.99+ 97-98 Nitrate 0.74 9.42 0.003 99.9690-95 Sulfate 46.1 278.4 <0.001 99.99+ 99-91 Boron 0.083 0.62 0.00798.51 60-70 (Dissolved) 67 480 0.46 99.87 95-99 Silica TOC 0.64 1.1<0.003 99.66+ 90-95 pH 8.0 10.8 10.3 — —Concentrations in ppm as ion, unless otherwise noted.

TABLE 2 Sodium Ion Exchange Effects Sodium, ppb Conventional New ROProcess RO Permeate 193 955 Post Cation IX 0.431 <0.007

TABLE 3 POST MIXED BED ION EXCHANGE RESULTS Conventional New ConstituentRO Process Boron Non-detectable Non-detectable Silica 0.43 ppb 0.35 ppbTOC  5.9 ppb <3.0 ppb

Specifically, the high rejection rates of boron and TOC also providesignificant additional benefit in reducing loading of downstream anion42 and mixed bed ion exchange units 44 and 46. In this regard, note thata rejection of 98.51% was achieved for boron, compared with about 60% to70% which is achievable in conventional RO systems on the samefeedwater. Typically, termination of an anion or mixed bed exchange runis determined by silica, or in certain cases, boron leakage. In spite ofhigher recovery in the pilot RO system, silica content in theconventional RO system permeate was three times higher than in the pilotRO system. Specifically, silica concentrations of 0.45 ppm SiO₂ wereachieved in permeate from the pilot test unit of this method, comparedto 1.5 ppm SiO₂ in conventional RO permeate. Clearly, levels of lessthan 1.0 ppm SiO₂ are achievable in RO permeate when utilizing thepresent method, and in fact, levels of less than 0.5 ppm SiO₂ have beenshown achievable. Also, the boron content in permeate from my novelprocess was 0.007 ppm B, versus 0.06 ppm B for permeate from aconventional RO system. Clearly, boron levels of less than 0.05 ppm weredemonstrated, as well as levels of less than 0.01 ppm of boron. The testresults from Table 1 also shown this result, in that rejection of boronin a conventional RO system ranges from about sixty percent to seventypercent (60%-70%), whereas rejection of boron in my water treatmentprocess was shown to be about ninety eight and one-half percent (98.5%).In other words, in a conventional RO process roughly thirty to forty (30to 40) borate ions pass through the membrane for each one hundred (100)present in the feedwater, whereas in my process less than two, andspecifically, only about one and one-half (1.5) borate ions pass throughthe membrane out of every one-hundred (100) present. In other words, 30per 100 or 40 per 100 borate ions in the feedwater reach the permeate inconventional RO, versus 1.5 per 100 in this process. In certainfeedwaters this number would decrease even further, to as low as 1/100,or 1/1000, for boron rejection rates of ninety nine percent (99%) orninety nine point nine percent, (99.9%), respectively. Thus, thisindicates that the run times on anion exchanger 42, while notnecessarily proportionate to the influent silica and boron levels, arenevertheless going to be significantly longer when treating permeate 34from my new process, as compared to run times when treating permeatefrom a conventional RO system. Since anion exhaustion is indicated by apredetermined level of leakage of silica (SiO₂), and, in some casesboron, and since the resin bed outlet concentration is related to themean species concentration in the resin bed, by achieving significantreduction in the concentration of such anions in the influent to theanion ion-exchange resin bed, the consequence is that longer run timesare attained before the maximum allowable leakage of SiO₂ or boron isreached.

Importantly, the levels of boron, and particularly silica and TOC werefound to be extremely low after treatment of the permeate 34 in themixed bed ion exchangers 44 and 46 in the pilot plant. A comparison withpost mixed bed permeate from a conventional RO process is provided withthe data in TABLE 3. Significantly, in my new process, in post mixed bedion-exchange treated water, the TOC level was found to be less than 3.0ppb, i.e., below detection limit.

And, not to be overlooked, are the significantly improved rejection ofsodium and potassium, which improved to 99.73% and 99.98%, respectively,from conventional RO system rejection rates ranging from ninety five toninety eight percent (95%-98%) in the case of sodium, and from aboutninety to ninety five percent (90%-95%), in the case of potassium.

The significantly higher rejection of strongly ionized species such assodium, potassium, chloride, and sulfate, compared to conventional ROoperations as evidenced by the data in Table 1, was a particularlyimportant and an unexpected experimental result of pilot testing.Further, even though the RO permeate in the pilot plant testingcontained a higher level of sodium than does the permeate of aconventional RO process, as noted in TABLE 2, the impact of the highersodium content on post RO cation exchange is relatively inconsequential.Since the RO permeate from my novel process is highly alkaline (atypical pH of 10.3 during pilot testing is shown in Table 1) andcontains significant levels of free hydroxide ions, the sodium removalextent, and capacity of the resin in cation exchange unit 40, isincreased by a substantial margin. The effect of the increased hydroxidealkalinity in the permeate to enhance removal of sodium from suchpermeate is shown in TABLE 2. In conventional RO treatment of the samefeedwater, where the RO system permeate has only 193 ppb of sodium, yetthe cation ion-exchange resin is only able to effect sodium removal toabout 0.431 ppb. In contrast, my novel process, even though 955 ppb ofsodium was encountered in the RO permeate after cation ion-exchangetreatment, the sodium ion concentration was reduced to less than 0.007ppb.

The improved rejection of total organic carbon (“TOC”) in my processalso provides a significant benefit to RO plant operators. It is normalfor waters of natural origin to contain detectable quantities of highmolecular weight organic acids and their derivatives, particularlyhumic, fulvic, and tannic acids. These compounds result from decay ofvegetative materials, and are usually related to condensation productsof phenol-like compounds. Broadly, humic acids include the fraction ofhumic substances which are soluble in water at alkaline pH, but whichprecipitate at acidic pH. Fulvic acids include the fraction of humicsubstances which are water soluble at alkaline and acidic pH. Theseacids, and their decomposition products, can be carried around in thefeedwater stream and form undesirable deposits on selected substrates,particularly anion selective substances. Also, they tend to contributeto fouling in conventional RO systems. Therefore, it is desirable tominimize the effect of such molecules on or through the reverse osmosismembrane, so that adverse consequences of their presence can be avoided,particularly at the anion ion-exchange unit. As can be seen by referenceto Table 1, the TOC content of the permeate 34 is substantially lower incomparison to TOC from a conventional RO process with identical TOC inthe raw feedwater. Specifically, there is rejection of ninety nine pointsixty six percent (99.66%) of TOC in the pilot plant RO system, comparedto only ninety to ninety five percent (90 to 95%) recovery inconventional RO systems. As in the cases of silica and boron, increasedionization of TOC at the elevated pH of my new process attributes tothis important result. Thus, taking advantage of the ionization range ofionizable organic carbon species enables effective TOC reductions whenoperating RO systems according to the method set forth herein.

Operational results of the pilot test unit may also be betterappreciated by reference to FIGS. 5, 6, 7, and 8. FIG. 5 illustrates therelationship between the axial differential pressure (ΔP) versus time,in pounds per square inch, for the reverse osmosis membrane employed inthe pilot reverse osmosis test equipment. The differential pressureshown has not been corrected for changes in feedwater flowrate. Incomparison to conventional RO, the pilot test results show that a stablenormalized permeate flow rate, a stable silica rejection rate, and astable differential pressure have been achieved. This indicates thatfouling/scaling have been essentially eliminated in my new process. FIG.6 shows the normalized permeate flow, in liters per minute, versus timeover a six month period, for the reverse osmosis membrane employed inthe pilot reverse osmosis test equipment.

FIG. 7 illustrates the silica concentration in the reverse osmosisreject stream over a six month period in pilot reverse osmosis testequipment. FIG. 8 illustrates the rejection percentage of silica, versustime over a six month period, for the reverse osmosis membrane employedin pilot reverse osmosis test equipment. This silica rejection is basedon an arithmetical mean silica concentration in the pilot RO unit.

After each shutdown of pilot plant operation due to a ten percent (10%)or more decline in normalized permeate flow, the membranes wereinspected and cleaned. An important finding was that cleaning can besimply and effectively accomplished by commodity membrane cleaningchemicals, such as hydrochloric acid solutions, tetrasodium EDTA, andsodium hydroxide. Expensive proprietary chemical cleaning agents werenot required. An RO membrane operated with feedwater pretreatment in themanner set forth herein was proven to be completely restored to a fluxof essentially one hundred percent (100%) of startup performance values.Substantially all of the cleaning was accomplished with the acidic firststep of the cleaning process, thus indicating that calcium carbonate,magnesium hydroxide, magnesium silicate, and the like, were thepredominant scaling species. Importantly, this revealed that neithersilica scaling or biofouling were major concerns under the specifiedprocess conditions. The enhanced runnability, or increased systemavailability, with minimal scaling and virtually non-existentbio-fouling, is clearly another important benefit of my novel ROoperational method.

Biological fouling of thin film composite membranes has heretoforetended to be a common problem, and, with certain specific feedwatersources, has been virtually insurmountable. Although it was anticipatedthat control of biological fouling would be improved due to operation atrelatively high pH levels, the degree of biological fouling controlactually achieved far exceeded expectations, with bacteria levels beingvirtually non-detectable during autopsy of RO membrane elements. Thismeans that instead of accumulating living and dead bacteria against themembrane surface, as is common in conventional RO systems, in my uniquemethod, incoming bacteria are killed and dissolved away from themembrane surface. Thus, this method of RO pretreatment and operation maybecome useful for treating problematic water sources. This is effectivebecause high pH solutions cause disinfection by cell lysing or ruptureof the cell wall. This is a quite potent and quick acting method ofanti-bacterial activity, when compared, for example, with chlorinationwhich acts by the much slower method of diffusion through the cell wallto cause death by inactivation of the microorganism's enzymes. Also incontrast to chlorine sanitized systems, at the high pH operationpreferred in the present method, viruses and endotoxins(lipopolysaccharide fragments derived from cell walls of Gram-negativebacteria) are effectively destroyed by lysis, thus enabling the presentmethod to be employable for the production of pyrogen free or sterilewater. In essence, the present method, when operated at a pH in excessof about 10, provides sanitization (3 log reduction in bacteria anddestruction of vegetative matter), and may also prove to essentiallyprovide true sterilization (12 log reduction in bacteria and theelimination of biofilm and spores) of the process equipment, as testresults showed a zero (0) bacteria count in the permeate. Also, itshould be noted that the increased pH of permeate in this method ofoperation enables similar, helpful results in the post RO treatmentequipment. Such a method of operation should be of particular benefit inthe production of high purity water for pharmaceutical applications,where the requirements for United States Pharmacopeia 23 (“USP 23”)standards, as supplemented, must ultimately be met by the final productwater. In this regard, the avoidance of use of raw water polymers,antiscalants, and other proprietary chemicals in RO pretreatment, asdescribed herein with respect to a preferred embodiment, can eliminateundesirable additives to pharmaceutical grade water, and reduce costs byreducing the necessary tests on RO product water. More concisely, theselection of a pH for RO operating conditions which does not supportbacteria growth, and carrying out of hardness and alkalinity removal toa level which avoid use of additives, is a superior method forproduction of high purity water.

A further benefit of high pH operation is with respect increasedprotection of membranes, particular the thin film composite types, whichhave limited tolerance for oxidizing agents at neutral, near neutral,and moderate alkaline pH's (up to roughly pH 9). When chlorine is addedto RO feedwater, gaseous chlorine (Cl₂) or sodium hypochlorite (NaOCl)are typically utilized. Because of membrane sensitivity to freechlorine, in conventional RO systems, it is normally removed by sulfite(SO₃ ⁻) injection. However, above pH 9, and particularly above pH 10,the effect of chlorine and other similar oxidants on thin film compositemembranes is significantly reduced. This is because the concentration ofthe non-ionized species (such as HOCl, known as hypochlorous acid) isdecreased dramatically, since such acids are relatively weak.Consequently, in my HERO™ high pH reverse osmosis process, typicallyoperating at a pH of 10 or higher, chlorine removal is not generallynecessary, thus reducing system complexity and costs. This may beespecially beneficial for those systems which utilize a municipal watersource as the feedwater to the water treatment plant.

Enhanced membrane life is also another benefit of my novel membraneoperation process. In membrane operations, and in particular withrespect to RO operations, longer membrane element life may be expected,primarily because scaling and biofouling are avoided, and thus, exposureto harsh cleaning chemicals (for instance, acid chemicals andsurfactants) is reduced dramatically.

RO membranes are taken out of service when the rejection of criticalspecies, for example silica, boron, or TOC, falls below an acceptablelimit. For silica, this usually occurs when rejection falls to betweenninety five and ninety six percent (95%-96%), from an original value ofninety nine percent (99%) or higher. As discussed above, the initialrejection values for silica in my process are significantly higher thanare achieved in conventional RO systems. Therefore, if conventional ROlimitations for silica rejection were accepted, for example, a specificmembrane element would last longer before the acceptable limits werereached. Stated another way, even after a considerable term of service,the membrane elements utilized in the present method will give silicarejections which are in excess of those provided by even new membranesoperating in conventional RO process configurations.

High flux, or permeate production, is also achievable due to the uniqueoperating conditions of my method for operating an RO system. Severalfactors contribute to this result. Flux, expressed as gallons of waterpassed through one square foot of membrane in one day, generally termed“GFD” (or, as liters of water passed through one square meter ofmembrane in one day, termed L/m²/day), is anticipated at about 15 GFDfor conventional RO systems. In pilot testing, the noted thin filmcomposite type FILMTEC BW membrane was operated at 24 GFD (977L/m²/day), and potential for up to 30 GFD (1222 L/m²/day) was favorablyevaluated. While the latter flux rate is believed to be the approximatecurrent hydraulic limit of conventional RO module design, based onspacer configurations, it is anticipated that even increased flux can beachieved in this method of operation (up to 50 GFD or so—which is 2036L/m²/day) when membrane modules become available that can support suchincreased flux. This is a most advantageous result for RO systemoperators, since, for example, if the normal flux is doubled by use ofthis method, then the total square feet (meters) of membrane surfacerequired is reduced by a factor of two. Corresponding decreases incapital cost (specifically, for membranes and pressure vessels) andfloor space requirements are therefore achieved. Operating cost, alreadysignificantly lowered by other benefits of the instant method, arefurther decreased by lowered membrane replacement costs. The one hundredfifty percent (150%) plus flux increase demonstrated in testing over thedesign basis for conventional RO systems provides an immediate benefit.

When utilizing the present method, osmotic pressure of the RO rejectrepresents the ultimate limitation for RO technology. Once appropriateraw feedwater treatment has effectively removed sparingly solublespecies, such as calcium carbonate, calcium sulfate, barium sulfate,silica, etc., then concentration of reject can proceed until the osmoticpressure limitation is reached. At this time, the design pressures forcommercially proven RO systems are typically limited to approximately1,200 psig (8280 kPa). If a design allowance is made for a 200 psigdriving force with respect to the reject stream, then the maximumallowable osmotic pressure would be approximately 1000 psig (6900 kPa).For purposes of example, based on a simplified rule of thumb thatapproximately one (1) psig of osmotic pressure is exerted by one hundred(100) ppm of TDS, the maximum allowable TDS of the reject stream wouldbe approximately 100,000 ppm. Thus, this new RO operating technology,regardless of feedwater chemistry, is potentially capable ofconcentrating any feedwater to approximately 100,000 ppm without concernwith respect to the various sparingly soluble species, and inparticular, with respect to calcium sulfate, barium sulfate, and silica.

Yet another advantage of my new RO operating technology is that existingRO systems, when retrofitted with the herein discussed pretreatmentequipment for hardness and alkalinity removal, can take advantage of theoperating benefits of this process method.

Additional applications for this unique RO operating method exist inboth high purity applications such as semiconductor manufacturing andpharmaceutical applications, as well as the more traditional industrialuses for boiler feedwater, cooling tower makeup water, and scrubbermakeup water. Application of my method of reverse osmosis systemoperation to high purity water production systems is shown in FIG. 9. Inthis figure, a multipass reverse osmosis process technique is utilizedto sequentially process a portion of initial raw feedwater 10 feedwaterto produce a final permeate 34 _((1+N)) which has been passedsequentially through a number N of reverse osmosis membrane units, whereN is a positive integer, typically two (2) or sometimes three (3),although a higher number could be utilized. As described above the rawfeed water 10, if deficient in alkalinity, may have alkalinity added byany convenient technique, such as by sodium carbonate 13, and then thattreated stream Rc is sent to the weak acid cation ion-exchange system12. After cation exchange, acid 14 such as hydrochloric or sulfuric maybe added to produce an intermediate treated stream W_(A). Then, carbondioxide is stripped in decarbonation unit 20 to produce an intermediatetreated stream D. Then, the pH is increased by a convenient and costeffective method such as addition of alkali solution 22 or by anionion-exchange unit 31, to produce a further intermediate treatment streamD_(OH). Reject 32 _((N)) from reverse osmosis unit N (and anyintermediate RO units between the first RO unit 30 ₍₁₎ and the final ROunit 30 _((N)) are then recycled into the feedwater before the RO unit30 _(N), to produce a feedwater 25 _((1+N)) containing undesirable buttolerable solute species and solvent water. Pump 26 pressurizes thefeedwater 25 _((1+N)) to produce a pressurized feed to the first RO unit30 ₍₁₎; after processing, permeate 34 ₍₁₎ results, which is then fed tothe next reverse osmosis unit in the series from 1 to N units. Thereject from the entire RO train is shown as reject 32 _((1+N)). Highpurity treated permeate from the entire train is shown as permeate orproduct water 34 _((1+N)), and it is fed to the usual ion-exchangeequipment for final cleanup before use. Cation ion-exchange unit 40produces a further intermediate purity stream C, which is followed byanion ion-exchange unit 42 to produce a further intermediate puritystream A. Before use, a primary mixed bed ion-exchange unit 44 producesa yet higher purity stream P, and an optional secondary or polishingmixed bed ion-exchange unit 46 produces a still higher purity, possiblefinal purity product S, or, using the same nomenclature as above, pureproduct water stream 38 _((1+N)). In semiconductor manufacturing, finalfiltration in sub-micron filters 48, using nominally sized 0.02 micronfilters, but perhaps selected from sizes ranging from about 0.02 micronto about 0.1 micron in size, is generally practiced, to produce a stillhigher product stream M. Also, biological control by passing high puritywater through a UV sterilizer unit 49 is customary, normally operatingat 254 nm wavelength to kill any bacteria which may remain in the highpurity stream M to produce a final ultrapure water U. In many systems,the positions of the final sub-micron filters 48 and the UV sterilizerunit 49 may be reversed, or a further post UV filter may be utilized.

FIG. 10 illustrates the use of my method of reverse osmosis systemoperation for boiler feed makeup water, or for cooling tower makeupwater, or for scrubber makeup water. The reverse osmosis unit 30 andvarious pretreatment equipment is operated according to the methods setforth hereinabove, to produce a high purity permeate 34. The productwater permeate 34 is then treated in an ion-exchange system as necessarybased on specific boiler requirements, and fed as makeup water 100 to aboiler 102. Blowdown 104 from boiler 102 is sent to an accumulation tank106 for pumping 108 through return line 109 to the RO pretreatmenttrain. Although the cooling tower 110 and scrubber 112 could be fed withRO permeate 34, more typically, the cooling tower 110 and scrubber 112,for example in a steam-electric power plant, would be supplied by usualraw water 10 supplies, such as municipal or well water. Therefore,cooling tower blowdown 114 and scrubber blowdown 116 are typically highin both hardness and alkalinity. Likewise, this system may be used totreat water having intimate contact with ash, such as ash pond water orash sluicing water from coal fired steam-electric power plants. In myreverse osmosis process, a significant amount of reusable water canusually be obtained by my method of RO pretreatment and operation,unlike the case with conventional RO systems.

Another advantageous use of my method for pretreatment and operation ofan RO system is illustrated in FIG. 11, where a preferred embodimentsimilar to that explained above is shown in use with a multipass ROsystem (here, two pass with RO units 30 ₍₁₎. and 30 _((N)), where N=2),in pretreatment for a continuous electrodeionization system 150. ROpermeate 34 _((1+N)), when treated by continuous electrodeionization,will produce a very high quality deionized water E which, afterultraviolet treatment 46 and final filtration 48, will be acceptable foruse in the microelectronics industry as ultrapure water UP. Optionally,the secondary or polish type mixed bed ion exchange unit 46 may beomitted, and the continuous electrodeionization product water E may besent directly to the UV sterilizing unit 49. This is true since thelimitations of continuous electrodeionization to reject boron, silica,TOC and the like thus limit its ability to produce, as a directeffluent, 18.2 megohm water for electronics manufacturing. Yet, thepermeate 34 _((1+N)) from the two pass RO system, when operatedaccording to the method disclosed herein, contains very low levels ofsuch species which are troublesome to continuous electrodeionization,such as boron, silica, TOC and the like. Thus, use of such permeate asfeed for a continuous electrodeionization treatment unit is believed toenable such electrodeionization units to produce 18.2 megohm waterwithout the benefit of downstream ion-exchange polishers 46. Theadvantage of using continuous electrodeionization over conventional ionexchange, of course, is that the continuous process (rather than thebatch process of ion exchange resins) is regenerated electrically,rather than chemically, and therefore avoids the use of conventionalregeneration chemicals.

And, even in wastewaters, the instant method may often be used toadvantage. Since an RO system when operated as taught herein willsubstantially reject ionizable species at high pH, high rejection ofsuch constituents will be achievable to produce an RO permeate low insuch constituents, for recycle and reuse. Wastewaters from refineries,pulping and papermaking operations, and municipal sewage treatmentplants, all are fairly high in candidate components (aliphatic and oraromatic organic acids and their derivatives), and are most difficultfor conventional RO membranes to handle due to organic fouling andrelated biological growth. Typical industrial uses where water ofsufficient quality may be attained when treating wastewaters includecooling towers, boiler makeup, scrubber makeup, and the like.

Benefits of HERO Brand RO Process Design and Operation

Many exemplary and desirable process benefits provided by the HERO brandRO system process design and operation were listed above at pages 15-16.Detailed explanation of such benefits include:

(A) High Rejection of Contaminants

As shown in Table 4, which summarizes data from a HERO brand RO processpilot plant, rejection of all species is significantly higher than whatcan be achieved in conventional RO operation. Particularly noticeable isthe improvement in the rejection of weak anions such as TOC, silica, andboron. Given that humic/fulvic acid derivatives (TOC), silicic acid, andboric acid are all relatively weak acids, at high operating pH theseacids will dissociate to a much greater extent (compared to near-neutralpH operation) and, therefore, will be much better rejected by the ROmembrane.

The improvement in the rejection of strongly ionized (at near-neutralpH) species was also observed. Several factors are believed tocontribute to the improvement in rejection of strongly ionized species.A change of membrane morphology, is believed to occur. A significantreduction in the thickness of the concentration polarization layeradjacent to the membrane surface (due to reduced surface tension at highfree causticity conditions) is believed to be a major contributor tothis improvement. Also, swelling of elastomers such as o-rings, and theresultant better sealing characteristics in the modules are also afactor.

The impact of much higher rejection of silica, etc., on thebehavior/operation of a post-RO ion exchange system is extremelysignificant. Since the vast majority of post-RO ion exchange isregenerated on the basis of either silica or boron breakthrough, afactor of ten reduction in the influent silica/boron content willprovide much longer run times between regenerations. Absence of carbondioxide, as well as bicarbonate in the RO permeate (due to a high pH,typically at least 10), will also increase on-line time beforesilica/boron leakage exceeds normal threshold values. Reduction ofstrongly ionized species concentration in the RO permeate is ofrelatively less significance, since most post-RO ion exchange isultimately silica or boron limited.

TABLE 4 COMPARISON OF HERO ™ RO VS. CONVENTIONAL RO Rejection RejectionPassage Passage Passage Passage (%) (%) (%) (%) Factor ReductionConstituent Conventional HERO Conventional HERO Conv/HERO (%) Sodium 9899.73 2 0.27 7.4 87 Potassium 90 99.98 10 0.02 500.0 99 Chloride 9899.99 2 0.01 200.0 99 Silica 99 99 99.87 1 0.13 7.7 87 Boron 70 98.51 301.49 20.1 95 TOC 95 99.66 5 0.34 14.7 93

TABLE 5 WATER ANALYSIS Raw Water RO Reject RO Product Na + K 125 1,350<1 Ca 7 0 0 Mg 13 0 0 HCO3 85 50 <1 CO3 0 50 <1 NO3 1 10 <1 SO4 30.8 308<1 Cl 28.2 282 <1 SiO2 50 500 <1 pH 7.1 10.8 10.2 Notes: 1. Analysis ofRO feed is not shown in the table, nor is the hydroxide content of ROreject and RO Product. 2. The chemistry is based on 90 percent ROrecovery, while maximum recovery feasible is approximately 96 percent.3. Except for pH, all constituents are reported as mg/l as CaCO₃

TABLE 6 COST ESTIMATE OF A RETROFIT Water/Waste Water Savings 244,000 US$/Yr) Antiscalant Elimination  30,000 US $/Yr) Power Savings  17,000 US$/Yr) Additional Chemical Costs (40,000) US $/Yr) AdditionalMiscellaneous Costs (20.000) US $/Yr) Net Annual Savings 231,000 US$/Yr) Conversion (Capital) Cost 200,000 (One Time) Simple Pay-BackPeriod     10.4 (Months)

TABLE 7 COST COMPARISON Conventional HERO System System EquipmentCapital Cost 12 7.8 (US $MM) Operating Cost 5.75 <4.00 (US $/1.000 USGallon) Note: See Section 5.0 for basis.

Compared to 60 to 70 percent boron rejection in conventional thin filmcomposite RO operation, the new process provides approximately 99percent boron rejection. In a double pass configuration, the new processis capable of producing a permeate with lower than detectable limits ofboron content.

Another very significant advantage in operating ion exchange withpermeate from a HERO brand RO system is that sodium leakage from cationresin is reduced by several orders of magnitude, due to the high ambientpH of the influent. As a result, longer run times between regenerationfor existing ion exchange systems means lower chemical and manpowerneeds, lower regeneration waste volume, etc. For new systems, or forexisting systems undergoing expansion, the new HERO brand RO processdesign and operation can have a strong positive impact on the ionexchange system capital cost as well.

(B) High Recovery Rates

Since hardness-causing ions such as calcium, magnesium, barium,strontium, aluminum, iron, manganese, etc., have been removed prior tothe RO, undesirable precipitation of species such as calcium carbonate,calcium fluoride, calcium sulfate, barium sulfate, magnesium hydroxide,aluminum/magnesium silicate, etc., does not occur in the HERO brand ROprocess, and thus that type of precipitation no longer limits therecovery achievable by an RO system. Importantly, silica solubility, isincreased dramatically at the normal HERO brand RO operating pH(preferably of approximately 11). Sustainable long-term operation withsilica levels in the 450 to 500 ppm range (in the RO reject) has beenproven, and theoretical models indicate that levels of 1,000 ppm orhigher may be achievable in this new RO operational method.

Based on 25 ppm silica in the RO feed, 95 percent recovery RO operation(approximately 500 ppm in the reject) has been proven by testing. Still,97.5 percent recovery (approximately 1,000 ppm silica in the RO reject)is theoretically feasible, whether or not practical from an operationalpoint of view. Since silica usually represents the ultimate limitingcriterion, in terms of maximum allowable recovery in an RO system,increased silica solubility along with essentially total absence ofspecies such as calcium, barium, etc., in RO feed, should allow ROoperation at very high recovery rates (90 to 98 percent) with the vastmajority of feedwaters.

With feedwater relatively high in barium content, RO system recovery canbe limited by barium sulfate precipitation potential at the reject end.The HERO system eliminates this concern altogether, since barium isquantitatively removed prior to the RO. The same outcome is alsoapplicable for RO systems limited (in recovery) by strontium sulfate,calcium sulfate, calcium fluoride, and other sparingly soluble calcium,magnesium, iron, and aluminum salts.

Of course, the final limit in RO recovery, represented by osmoticpressure of the RO reject, will still control the maximum feasiblerecovery achievable with a specific feedwater, but this limit is notusually reached at recoveries less than 99 percent with most feedwaters.

(C) Biological Fouling is Essentially Eliminated

Most commonly occurring microbial species are completely lysed(physically destroyed) at the high operating pH. In fact, even virus,spores, and endotoxins are either destroyed or rendered incapable ofreproduction/proliferation at very high pH levels. Saponification oflipids (fat) is expected to play a role in the process as well sincefatty acids and their corresponding glycerides will form soluble “soaps”at the high operating pH.

In one location where long-term tests were carried out, biofouling wasconspicuous by its absence during the test of the HERO technology. Thispilot RO system exhibited very stable operating performance in terms ofnormalized permeate flow and system pressure drop throughout the testperiod. Further confirmation of the absence of biofouling was obtainedduring autopsy of RO elements at regular intervals. A stage wise programto test and autopsy the FILMTEC FT30 based elements was conducted over a15 month period. The data showed higher salt rejection than the initialQuality Assurance values under standard test conditions. Also, themembrane surface was clean and free of any evidence of biofouling.

This characteristic of the new process can be of significant benefit forsites with known biofouling problems or for the treatment ofbio-contaminated/bio-active wastewater. It can also be very effectivefor systems with higher-than-ambient temperature RO operation.

(D) Particulate Fouling is Substantially Reduced

It has been known (and practiced) for almost 30 years that softening ofRO feedwater destabilizes colloidal solids present in the feedwater andsignificantly reduces the associated fouling problems. Mandatorysoftening requirement as pretreatment for hollow fine fiber RO elementsin the late 1960s and early 1970s attests to this strategy. In addition,zeta potential is generally reduced between a surface and foulantparticles at high pH, thus reducing the likelihood of adhesion. Thisproperty is accentuated by the fact that most naturally occurringparticles (including bacteria) exhibit negative surface charges. Whileside-by-side zeta potential determination is yet to be carried out, thenew process is expected to significantly reduce, if not eliminate,particulate fouling problems. The reduction of zeta potential furtherreduces the possibility of particle adhesion to the slightly negativelycharged-membrane surface. The in-situ formation of surfactants frombacterial lipids, if present, will further help in reducing particleadhesion to the membrane surface.

This unique characteristic of the new process can be of significantvalue in the design of an RO system, particularly in the potential toreduce capital cost and operating complexity of treating UPW. Inaddition to the ability to accept a certain level of particulatefoulants, the new process may also minimize the need for multimediafiltration, coagulant/floucculant addition, Diatomaceous Earthfiltration, etc., as pretreatment to the RO system.

(E) Significantly Reduction in Chemical Usage

Dechlorination, either by chemical addition or by activated carbon, mayvery well be unnecessary as well since the level of free (undissociated)hypochlorous acid (HOCL) is extremely low at the very high operating pH.

(F) Elimination of Scale Inhibitor Use

Use of antiscalants or scale inhibitors, while not harmful orincompatible with the new process, can be completely eliminated, asproven by an 18-month test at a semiconductor manufacturing facility.

(G) High Flux Rates

Given the reduced thickness of the concentration polarization layer, aswell as the elimination of biofouling and the reduction of particulateadhesion to the membrane surface, it is not surprising that an RO systemutilizing the new process can operate at higher flux compared toconventional operation. Compared to a normal design flux of 15 gallonsper square foot per day (GFD) (611 L/m²/day), the HERO brand RO systemis designed in excess of 15 GFD (611 L/m²/day), and is preferablydesigned at about 20 GFD (814 L/m²/day), and more preferably up to about25 GFD (1018 L/m²/day), and, where feasible, in excess of 25 GFD (1018L/m²/day).

(H) Higher Product Purity

In addition to reduced capital cost for the RO system, the quality ofthe RO permeate is improved significantly due to the higher design flux.For example, at 25 GFD, the RO permeate will contain 40 percent lowerdissolved solids compared to a 15 GFD (611 L/m²/day) design basis. Thehigher pH operation, in combination with the high product flux, providesthe result that the salt flux (which is concentration dependent, ratherthan pressure dependent) is significantly reduced. The RO system can beexpected to be about 20 percent less expensive due to this factor alone(or more than 20 percent less expensive), all other parameters beingequal.

(I) Reject Usable as Scrubber Makeup

The reject from the HERO brand RO system, with high pH, low carbonatealkalinity, and virtually no hardness, can be used as makeup to acidicgas scrubbers. Due to concerns about potential silica precipitation ifthe pH is lowered significantly in the scrubber, the RO reject should beused on a once-through basis, and thus not be evaporation rate limited.

Process Chemistry

As discussed earlier, very high reject pH is one factor whichcharacterize operation of the HERO brand RO system. Extremely highrejection of the weakly ionized anions such as TOC, silica, boron, etc.,can be correlated to such characteristics. The following example, basedon silica, can be used to explore this relationship.

In naturally occurring waters and at near-neutral pH range (6-8), silicais primarily present as orthosilicic acid (H₄SiO₄). Orthosilicic acid,commonly referred to as silicic acid, is one of the weakest acid speciespresent in water. Silicic acid's first dissociation constant (i.e. thedissociation of the first proton from the total of four hydrogens) isapproximately 2×10⁻¹⁰, corresponding to a pKa value of approximately 9.7at ambient temperature and very low background ionic strength of thesolution.

A convenient way of visualizing the relative strength of silicic acidwith pKa₁ of 9.7 is to state that at pH 9.7, it is fifty percent (50%)percent ionized, i.e. 50 percent of it is present as undissociatedorthosilicic acid, while the other 50 percent is dissociated and ispresent as monovalent silicate ion, the conjugate base of orthosilicicacid. At pH 10.7, when the log of conjugate base to undissociated acidis unity, approximately 91 percent exists as silicate ion, the other 9percent as undissociated acid. At pH 11.7, the distribution is 99 and 1percent respectively. Conversely, at pH of 8.7 (when log of ratio is0.1), approximately 91 percent of the species is present asundissociated acid and 9 percent as the ionized silicate. At a pH of7.7, approximately 99 percent is present as undissociated silicic acidand 1 percent as the ionized monovalent silicate ion.

Since the majority of naturally occurring feedwaters are at pH 8 orlower, essentially all the silica exists as undissociated silicic acidunder these conditions. Other very weak acids, such as boric acid(H₃BO₃, with pKa, of approximately 9.2) and hydrocyanic acid (HCN, withpKa of approximately 9.3) exhibit very similar properties, but of coursethey are both somewhat stronger acids compared to silica.

Rejection characteristics of individual species across the RO membraneis influenced by the size, shape, and charge density of the solute. Itis generally recognized that an ionized solute will be rejected mustbetter compared to a solute that exists in an undissociated state,provided that their size and shape are comparable. Rejection offluoride, for example, is essentially zero at pH less than 3, percent atpH 3.5, 50 percent at pH 4, 75 percent at pH 5, and 98 percent (or more)at pH 7. Hydrofluoric acid (a weak acid with pKa of 3.2) is thecounterpart of the ionized fluoride species and is the primary componentat low pH values.

Rejection of silica/silicic acid, however, is a surprisingly high 98percent at pH 7, where the primary constituent is the undissociatedsilicic acid and not the ionized silicate species. This discrepancy isat least partially explained by the fact that the actual size of (ortho)silicic acid is much bigger than expected since the moleculeincorporates up to six molecules of water of hydration. Thus, the highrejection is due to the size/shape factor, since at pH 7 there is verylittle ionization (less than 0.2 percent) of silicic acid.

Based on the factors involved, it would appear that silica, whensubstantially ionized, should have rejection comparable to that ofsulfate (SO₄ ⁼) ion. The expectation is based on the fact that thesulfate ion also incorporates six waters of hydration and, of course, itis completely ionized at near-neutral pH values. As a matter of record,sulfate rejection of 99.5 to 99.9 is routinely observed in normal ROoperation and the silica rejection in the HERO system operating at pH10.5 to 11.0 range has actually been better than 99.9 percent. In otherwords, sulfate rejection at pH 7 and silica rejection at pH above 10 arequite comparable. In view of the relative strengths of the correspondingacids and the relative size of the molecules, this effect can berationalized as well as utilized.

Another aspect of the new process that merits further discussion is therequirement for essentially complete removal of alkalinity prior to pHadjustment (increase) of the RO feed. From an entirely practical pointof view, near-zero alkalinity is a necessity since any residualalkalinity will provide a strong buffering effect and substantiallyincrease the quantity of alkali needed to raise the pH to the normaloperating range. Over and above the direct cost of increased alkalirequirement, the sodium content of the RO permeate will be much higheralso, resulting in unnecessarily high post-RO ion exchange load andcost.

From a conceptual point of view, however, the requirement for alkalinityremoval is far more urgent but straightforward. The following example,based on calcium carbonate solubility, will be used to quantify therelationship.

Solubility product (Ksp) of calcium carbonate is approximately 8.7×10⁻⁹square molar at ambient temperature and very low ionic strength.Assuming 90 percent recovery across the RO is the goal, allowablemaximum CaCO₃ ion product of the RO feed is approximately 8.7×10⁻¹¹square molar. Further assuming 0.1 mg/l of calcium in the softenedfeedwater, the allowable maximum carbonate content of the RO feed isapproximately 2.1 mg/l, all expressed as ions.

At pH 11.0 reject condition, approximately 85 percent of thecarbonate(s) species is present as carbonate, the rest exists asbicarbonate. Assuming 5 mg/l of total residual carbon dioxide equivalentprior to pH increase, approximately 5.8 mg/l of carbonate (as ion) willbe present in the RO feed. Compared to the maximum allowable 2.1 mg/l ofcarbonate, the achievable 5.8 mg/l is three times as high.

To ensure scale-free operation at 90 percent recovery, one or more ofthe following must be achieved—residual calcium content must be lessthan 0.1 mg/l, or the RO operating conditions must be changed. Whilecalcium carbonate scale inhibitors are known to generally allow a highKsp, I am not aware of any such formulation which would efficiently andcost effectively allow continuous high pH operation of RO. Important, itshould be noted that during the long-term testing of the HERO system, noscale inhibitors were used whatsoever.

Magnesium hydroxide, with a Ksp of approximately 1.2×10⁻¹¹ cubic molar,is in some ways even more demanding in terms of allowable residuals,since magnesium tends to leak earlier from the weak acid cationexchanger and, therefore, more care is needed to prevent magnesiumhydroxide scale.

Typical Example

The following is an example for a typical application for the HEROsystem. The feedwater in the Kumamoto area in Southern Japan, high insilica content, was selected for the example. Costs shown are budgetary(+ or −30 percent accuracy). A cost projection is based on the followingassumptions:

-   -   (1) 1,500,000 US Gallons Per Day (5,700 m³/day) system nominal        capacity;    -   (2) 75 percent normal recovery rate        -   vs.        -   90 percent HERO system recovery rate;    -   (3) UPW quality (chemical) criteria are:        -   (a) silica <1 PPB,        -   (b) TOC <1 PPB, and        -   (c) Oxygen <5 PPB;    -   (4) Consumable costs:        -   (a) Sulfuric acid (93 percent) at US$100/ton (US$110/metric            ton);        -   (b) sodium hydroxide (100%) at US$450/ton ($500/metric ton);        -   (c) antiscalant at US$1.50/pound (US$3.30/kg);        -   (d) electricity at US$0.075/kwh;        -   (e) water purchase and wastewater discharge costs (combined)            at US$3/1,000 gallons (US$0.79/m³).

Conversion of Existing RO System

Table 6 below assumes that an existing 1.5 million US Gallons Per Day(5,700 m³ day) system operating at 75 percent recovery with feedwatershown in Table 5 is converted into a 90 percent recovery HERO brand ROprocess system and that no changes are made beyond the RO system.

In some cases, it may also be feasible to use a HERO brand RO processdesign to increase overall RO recovery rates, by processing reject froma conventional RO system, by (a) simultaneously reducing hardness andalkalinity in a WAC system, (b) decarbonation, and (c) raising the pH,before feeding the stream into a second RO system.

Conversion of existing systems may also provide unique opportunities toincrease the capacity of an RO system. This is possible because the fluxof about 15 GFD (611 L/m²/day) in a conventional RO system can beincreased up to about 20 GFD (814 L/m²/day), or perhaps up to as much as25 GFD (1014 L/m²/day), or more, when the operation is changed to a HERObrand RO process design and operation configuration.

New RO System Design and Operation

The projection in Table 7 below is made on the basis that two brand newUPW systems will be built, in one case utilizing the conventionalapproach (see FIG. 13), and in the other case utilizing the HERO brandRO system (see FIG. 14) that includes a simplified polishing loopdesign. Both systems will use double-pass RO, hollow fine fiber ultrafilter and no dual-bed ion exchangers. Approximately 40 percent of theUPW usage will be at high temperature, and the cost estimate includesDIW heaters. The distribution piping system beyond the ultra filtrationsystem is not included in these cost estimates, nor is systeminstallation or any PVDF lined storage tanks, since sizing of thesecomponents are very site specific.

Summary

The new HERO brand RO technology has been shown to exhibit very highrejection of all contaminants, especially weak acid anions. In addition,RO recovery of ninety percent (90%) or higher can be achieved with thevast majority of feedwater. Biological fouling is essentially eliminatedwhile particulate fouling is substantially reduced. A flux considerablyhigher than is normally practical using conventional RO system designcan be achieved with the new HERO technology. Although the benefits ofthis new process might justify higher UPW system cost, just the oppositeis true. The overall cost as well as the complexity of the UPW systemare both reduced dramatically.

The method and apparatus for processing water via membrane separationequipment, and in particular, via the HERO brand reverse osmosis (“RO”)process design as described herein, provides a revolutionary,paradoxical result, namely, simultaneous increase in levels of silica inthe RO reject, but with lower levels of silica in the purified ROpermeate. This method of operating membrane separation systems, and inparticular, for operating reverse osmosis systems, represents asignificant option for reducing water use while simultaneously reducingcapital and operating costs of the water treatment system. Waterrecovery, that is, the ratio of the quantity of the permeate productstream produced to the quantity of the feedwater stream provided isclearly in excess of about 50%, and easily will be up to about 85% ormore, and often, will be up to about 95%, and, at times, will reachlevels of about 99%. Further, given the efficiencies, dramatically lessusage of chemical reagents, either for ion exchange regenerant or for ROcleaning, will be consumed per gallon (or liter) of pure water produced.

It will thus be seen that the objects set forth above, including thosemade apparent from the proceeding description, are efficiently attained,and, since certain changes may be made in carrying out the above methodand in construction of a suitable apparatus in which to practice themethod and in which to produce the desired product as set forth herein,it is to be understood that the invention may be embodied in otherspecific forms without departing from the spirit or essentialcharacteristics thereof. For example, while I have set forth anexemplary design for simultaneous hardness and alkalinity removal, otherembodiments are also feasible to attain the result of the principles ofthe method disclosed herein. Therefore, it will be understood that theforegoing description of representative embodiments of the inventionhave been presented only for purposes of illustration and for providingan understanding of the invention, and it is not intended to beexhaustive or restrictive, or to limit the invention to the preciseforms disclosed. On the contrary, the intention is to cover allmodifications, equivalents, and alternatives falling within the spiritand scope of the invention as expressed in the appended claims. As such,the claims are intended to cover the methods and structures describedtherein, and not only the equivalents or structural equivalents thereof,but also equivalent structures or methods. Thus, the scope of theinvention, as indicated by the appended claims, is intended to includevariations from the embodiments provided which are neverthelessdescribed by the broad meaning and range properly afforded to thelanguage of the claims, or to the equivalents thereof.

What is claimed is:
 1. A process for treatment of a feedwater streamwith reverse osmosis membrane separation equipment, to produce a lowsolute containing product stream and a high solute containing rejectstream, said process comprising: (a) providing a feedwater streamcontaining solutes therein, said solutes comprising (i) hardness, (ii)alkalinity, and (iii) at least one molecular species which is sparinglyionized when in neutral or near neutral pH aqueous solution, said atleast one molecular species comprising one or more of (1) at least someTOC, or (2) at least some silica, or (3) at least some boron; (b)effectively eliminating the tendency of said feedwater to form scalewhen said feedwater is concentrated at a selected pH, by treating thefeedwater stream in a hydrogen form cation exchange resin to (i) removesubstantially all hardness from said feedwater stream, (ii) removesubstantially all alkalinity from said feedwater stream, and (iii)remove dissolved carbon dioxide created during removal of said hardnessand said alkalinity; (c) raising the pH of the product from step (b) toa selected pH of ten (10) or more, by adding a selected alkali thereto,to urge said at least one molecular species which is sparingly ionizedwhen in neutral or near neutral pH aqueous solution toward increasedionization; and (d) passing the product from step (c) above throughreverse osmosis membrane separation equipment, to concentrate saidfeedwater at a recovery rate of ninety percent (90%) or more.
 2. Theprocess as set forth in claim 1, wherein at least some alkali is addedto said feedwater when the hardness of the feedwater is higher than thealkalinity thereof.
 3. A process for treatment of a feedwater streamwith reverse osmosis membrane separation equipment, to produce a lowsolute containing product stream and a high solute containing rejectstream, said process comprising: (a) providing a feedwater streamcontaining solutes therein, said solutes comprising (i) hardness, (ii)alkalinity, and (iii) at least one molecular species which is sparinglyionized when in neutral or near neutral pH aqueous solution, said atleast one molecular species comprising one or more of (1) at least someTOC, or (2) at least some silica, or (3) at least some boron; (b)effectively eliminating the tendency of said feedwater to form scalewhen said feedwater is concentrated at a selected pH, by treating (i)the feedwater stream in a sodium form cation exchange resin to removesubstantially all hardness from said feedwater stream, (ii) adding acidto the feedwater treated in step (b)(i), to remove substantially allalkalinity, and (iii) removing dissolved carbon dioxide created duringstep (b)(ii); (c) raising the pH of the product from step (b) to aselected pH of ten (10) or more, by adding a selected alkali thereto, tourge said at least one molecular species which is sparingly ionized whenin neutral or near neutral pH aqueous solution toward increasedionization; and (d) passing the product from step (c) above throughreverse osmosis membrane separation equipment, to concentrate saidfeedwater at a recovery rate of ninety percent (90%) or more.
 4. Aprocess for treatment of a feedwater stream with reverse osmosismembrane separation equipment, to produce a low solute containingproduct stream and a high solute containing reject stream, said processcomprising: (a) providing a feedwater stream containing solutes therein,said solutes comprising (i) hardness, (ii) alkalinity, and (iii) atleast one molecular species which is sparingly ionized when in neutralor near neutral pH aqueous solution, said at least one molecular speciescomprising one or more of (1) at least some TOC, or (2) at least somesilica, or (3) at least some boron; (b) effectively eliminating thetendency of said feedwater to form scale when said feedwater isconcentrated at a selected pH, by treating (i) the feedwater stream in acation exchange resin to remove an effective amount of hardness fromsaid feedwater stream, (ii) adding acid to the feedwater treated in step(b)(i), to remove an effective amount of alkalinity, and (iii) removingdissolved carbon dioxide created during step (b)(ii); (c) raising the pHof the product from step (b) to a selected pH of ten (10) or more, byadding a selected alkali thereto, to urge said at least one molecularspecies which is sparingly ionized when in neutral or near neutral pHaqueous solution toward increased ionization; and (d) passing theproduct from step (c) above through reverse osmosis membrane separationequipment, to concentrate said feedwater at a recovery rate of ninetypercent (90%) or more.
 5. A process for treatment of a feedwater streamwith reverse osmosis membrane separation equipment, to produce a lowsolute containing product stream and a high solute containing rejectstream, said process comprising: (a) providing a feedwater streamcontaining solutes therein, said solutes comprising (i) hardness, (ii)alkalinity, and (iii) at least one molecular species which is sparinglyionized when in neutral or near neutral pH aqueous solution, said atleast one molecular species comprising one or more of (1) at least someTOC, or (2) at least some silica, or (3) at least some boron; (b)effectively eliminating the tendency of said feedwater to form scalewhen said feedwater is concentrated at a selected pH, by (i) treatingthe feedwater stream in a cation exchange resin to remove substantiallyall hardness from said feedwater stream, (ii) adding acid to thefeedwater treated in step (b)(i) to remove substantially all remainingalkalinity, and (iii) removing dissolved carbon dioxide created; (c)raising the pH of the product from step (b) to a selected pH of ten (10)or more, by adding a selected alkali thereto, to urge said at least onemolecular species which is sparingly ionized when in neutral or nearneutral pH aqueous solution toward increased ionization; and (d) passingthe product from step (c) above through reverse osmosis membraneseparation equipment, to concentrate said feedwater at a recovery rateof ninety percent (90%) or more.
 6. The process as set forth in claim 3,or in claim 5, wherein said alkalinity comprises non-hydroxidealkalinity.
 7. The process as set forth in claim 1, or in claim 3, or inclaim 4, or in claim 5, wherein raising the pH of the stream to pH toten (10) or more is provided by an anion-exchange system.
 8. The processin claim 1, or in claim 3, or in claim 4, or in claim 5, wherein saidreverse osmosis membrane separation equipment comprises two reverseosmosis units operated in series with respect to said product stream. 9.The process as set forth in claim 1, or in claim 3, or in claim 4, or inclaim 5, wherein said reverse osmosis membrane equipment is operated ata water flux of more than 0.611 m³/m²/day (15 gal/ft²/day).
 10. Theprocess as set forth in claim 1, or in claim 3, or in claim 4, or inclaim 5, wherein said reverse osmosis membrane equipment is operated ata water flux of between 0.611 m³/m²/day (15 gal/ft²/day) and 1.222m³/m²/day (30 gallons per square foot per day).
 11. The process as setforth in claim 1, or in claim 3, or in claim 4, or in claim 5, whereinsaid reverse osmosis membrane separation equipment generates a productwater stream low in solute content and a reject stream high in a solutecontent, and concurrently increases product production and product waterpurity.
 12. The process as set forth in claim 1, or in claim 3, or inclaim 4, or in claim 5, wherein operation of said reverse osmosismembrane equipment is practiced continuously.
 13. The process as setforth in claim 12, wherein said process is further characterized by astable normalized product stream production rate.
 14. The process as setforth in claim 5, wherein at least some alkali is added to saidfeedwater when the hardness of the feedwater is higher than thealkalinity thereof.
 15. In a process for purification of an aqueoussolution comprising solutes and solvent by using membrane separationequipment, said membrane separation equipment comprising asemi-permeable reverse osmosis membrane that substantially resist thepassage of dissolved species therethrough, to increase the concentrationof said solution to a preselected concentration factor by passing saidaqueous solution through said semi-permeable membrane in a first unit ofsaid membrane separation equipment to produce a product stream and areject stream, the improvement which comprises (a) minimizing hardnessof said aqueous feed solution, (b) minimizing alkalinity of said aqueousfeed solution, and (c) minimizing carbon dioxide dissolved or suspendedtherein, before increasing the pH of said aqueous feed solution to ten(10) or more and then concentrating said feedwater at a recovery rate ofninety percent (90%) or more, by controlling solutes, includinghardness, alkalinity, and carbon dioxide in said aqueous solution to alevel where the tendency to form scale is effectively eliminated at saidpreselected concentration factor prior to feed of said aqueous solutionto said first unit of said membrane separation equipment.
 16. Theprocess as set forth in claim 15, wherein the process of minimizingalkalinity comprises removing substantially all non-hydroxidealkalinity.
 17. The process as set forth in claim 15, wherein theprocess of minimizing alkalinity comprises removing substantially allalkalinity associated with hardness.
 18. The process as set forth inclaim 1, or in claim 3, or in claim 4, or in claim 5, or in claim 15,wherein above feedwater is concentrated at a recovery rate of ninetyfive percent (95%) or more.
 19. A process for treatment of a feedwaterstream with reverse osmosis membrane separation equipment, to produce alow solute containing product stream and a high solute containing rejectstream, said process comprising: (a) providing a feedwater streamcontaining solutes therein, said solutes comprising (i) hardness, (ii)alkalinity, and (iii) at least one molecular species which is sparinglyionized when in neutral or near neutral pH aqueous solution, said atleast one molecular species comprising one or more of (1) at least someTOC, or (2) at least some silica, or (3) at least some boron; (b)effectively eliminating the tendency of said feedwater to form scalewhen said feedwater is concentrated at a selected recovery rate at aselected pH, by (i) treating the feedwater stream in a softener and thenby strong acid cation exchange resin to remove substantially allhardness from said feedwater stream, (ii) removing alkalinity from saidfeedwater stream, and (iii) removing carbon dioxide from said feedwaterstream; (c) raising the pH of the product from step (b) to a selected pHof ten (10) or more, by adding a selected alkali thereto, to urge saidat least one molecular species which is sparingly ionized when inneutral or near neutral pH aqueous solution toward increased ionization;and (d) passing the product from step (c) above through reverse osmosismembrane separation equipment, to concentrate said feedwater at arecovery rate of ninety percent (90%) or more.
 20. The process as setforth in claim 19, wherein said strong acid cation exchange system isoperated in the hydrogen form.
 21. The process as set forth in claim 19,wherein said strong acid cation exchange system is operated in thesodium form.
 22. A process for treatment of a feedwater stream withreverse osmosis membrane separation equipment, to produce a low solutecontaining product stream and a high solute containing reject stream,said process comprising: (a) providing a feedwater stream containingsolutes therein, said solutes comprising (i) hardness, (ii) alkalinity,and (iii) at least one molecular species which is sparingly ionized whenin neutral or near neutral pH aqueous solution, said at least onemolecular species comprising one or more of (1) at least some TOC, or(2) at least some silica, or (3) at least some boron; (b) effectivelyeliminating the tendency of said feedwater to form scale when saidfeedwater is concentrated at a selected recovery rate at a selected pH,by (i) treating the feedwater stream in a softener and then by weak acidcation exchange resin to remove substantially all hardness from saidfeedwater stream, (ii) removing alkalinity from said feedwater stream,and (iii) removing carbon dioxide from said feedwater stream; (c)raising the pH of the product from step (b) to a selected pH of ten (10)or more, by adding a selected alkali thereto, to urge said at least onemolecular species which is sparingly ionized when in neutral or nearneutral pH aqueous solution toward increased ionization; and (d) passingthe product from step (c) above through reverse osmosis membraneseparation equipment, to concentrate said feedwater at a recovery rateof ninety percent (90%) or more.
 23. The process as set forth in claim19 or in claim 22, wherein said softener comprises a lime softener. 24.The process as set forth in claim 19 or in claim 22, wherein saidsoftener comprises a lime/soda softener.
 25. The process as set forth inclaim 22, wherein said weak acid cation exchange system is operated inthe hydrogen form.
 26. The process as set forth in claim 22, whereinsaid weak acid cation exchange system is operated in the sodium form.27. The process as set forth in claim 19, or in claim 22, whereinremoving alkalinity from said feedwater stream comprises adding acid.28. The process as set forth in claim 19, or in claim 22, wherein saidreverse osmosis membrane equipment is operated at a water flux ofbetween 0.611 m³/m²/day (15 gal/ft²/day) and 1.222 m³/m²/day (30 gallonsper square foot per day).
 29. The process as set forth in claim 19, orin claim 22, wherein operation of said reverse osmosis membraneequipment is practiced continuously.
 30. The process as set forth inclaim 19, or in claim 22, wherein effectively eliminating the tendencyof said feedwater to form scale when said feedwater is concentrated at aselected recovery rate at a selected pH, is performed on the feedwaterin the order of step(b)(ii), then step (b)(iii), and finally step(b)(i).
 31. The process as set forth in claim 19, or in claim 22,wherein said reverse osmosis membrane separation equipment generates aproduct water stream low in solute content and a reject stream high in asolute content, and concurrently increases product production andproduct water purity.
 32. A process for treatment of a feedwater streamwith reverse osmosis membrane separation equipment, to produce a lowsolute containing product stream and a high solute containing rejectstream, said process comprising: (a) providing a feedwater streamcontaining solutes therein, said solutes comprising (i) hardness, (ii)alkalinity, and (iii) at least one molecular species which is sparinglyionized when in neutral or near neutral pH aqueous solution, said atleast one molecular species comprising one or more of (1) at least someTOC, or (2) at least some silica, or (3) at least some boron; (b)effectively eliminating the tendency of said feedwater to form scalewhen said feedwater is concentrated at a selected recovery rate at aselected pH, by (i) treating the feedwater stream in a softener and thenby strong acid cation exchange resin to remove substantially allhardness from said feedwater stream, (ii) to the extent required toeffectively eliminate the tendency of said feedwater to form scale whenconcentrated to said selected recovery rate at a selected pH, removingan effective amount of alkalinity from said feedwater stream, and (iii)removing carbon dioxide from said feedwater stream; (c) raising the pHof the product from step (b) to a selected pH of ten (10) or more, byadding a selected alkali thereto, to urge said at least one molecularspecies which is sparingly ionized when in neutral or near neutral pHaqueous solution toward increased ionization; and (d) passing theproduct from step (c) above through reverse osmosis membrane separationequipment, to concentrate said feedwater at a recovery rate of ninetypercent (90%) or more.
 33. A process for treatment of a feedwater streamwith reverse osmosis membrane separation equipment, to produce a lowsolute containing product stream and a high solute containing rejectstream, said process comprising: (a) providing a feedwater streamcontaining solutes therein, said solutes comprising (i) hardness, (ii)alkalinity, and (iii) at least one molecular species which is sparinglyionized when in neutral or near neutral pH aqueous solution, said atleast one molecular species comprising one or more of (1) at least someTOC, or (2) at least some silica, or (3) at least some boron; (b)effectively eliminating the tendency of said feedwater to form scalewhen said feedwater is concentrated at a selected recovery rate at aselected pH, by (i) treating the feedwater stream in a softener and thenby weak acid cation exchange resin to remove substantially all hardnessfrom said feedwater stream, (ii) to the extent required to effectivelyeliminate the tendency of said feedwater to form scale when concentratedto said selected recovery rate at a selected pH, removing an effectiveamount of alkalinity from said feedwater stream, and (iii) removingcarbon dioxide from said feedwater stream; (c) raising the pH of theproduct from step (b) to a selected pH of ten (10) or more, by adding aselected alkali thereto, to urge said at least one molecular specieswhich is sparingly ionized when in neutral or near neutral pH aqueoussolution toward increased ionization; and (d) passing the product fromstep (c) above through reverse osmosis membrane separation equipment, toconcentrate said feedwater at a recovery rate of ninety percent (90%) ormore.